Microchannel reactors for catalytic oxidative dehydrogenation

ABSTRACT

The invention provides methods of oxidative dehydrogenation (ODH). Conducting ODH in microchannels has unexpectedly been found to yield superior performance when compared to the same reactions at the same conditions in larger reactors. ODH methods employing a Mo—V—Mg—O catalyst is also described. Microchannel apparatus for conducting ODH is also disclosed.

RELATED APPLICATIONS

This application is a divisional of Ser. No. 10/441,921 filed May 19,2003, now U.S. Pat. No. 7,402,719, incorporated herein by reference asif reproduced in full below. In accordance with 35 U.S.C. sect. 119(e),this application claims priority to U.S. Provisional Application Nos.60/388,635, filed Jun. 13, 2002, which is incorporated herein as ifreproduced in full below.

FIELD OF THE INVENTION

The present invention relates to the production of olefinic hydrocarbonsby oxidative dehydrogenation. More particularly, this invention relatesto catalytic oxidative dehydrogenation of hydrocarbons to yield olefins,and preferably, to the production of light olefins from lighthydrocarbons and to the production of styrene from ethylbenzene.

BACKGROUND OF THE INVENTION

Olefinic hydrocarbons, such as ethylene, propene, butene, and isobutene,are critical intermediates in the petrochemical industry. In order tosatisfy market demand, substantial efforts have been invested in theproduction of such compounds by conventional thermal steam cracking ofalkanes and naphtha and by catalytic dehydrogenation methods. However,conventional steam cracking is equilibrium limited and requires veryhigh temperatures (over 700° C.) to achieve a high enough conversion ofethane to be economically viable. Even so, temperatures are limited byavailable alloys to temperatures at which single pass yields are stillrelatively low. Steam cracking also requires the input of large amountsof heat to drive the endothermic gas phase cracking reaction. Because ofthe equilibrium limitations, steam cracking must be carried out at lowpressures typically 1 atmosphere or less and requires cooling andcompression of the product stream to effect separation and recovery ofthe olefins produced.

Conventional catalytic dehydrogenation has similar disadvantages,including the need for high reaction temperatures (e.g., 550 to over700° C. depending on the feedstock), the deactivation of the catalyst bycoke formation, and the consequent need for continuous or periodiccatalyst regeneration at frequent intervals throughout the process. Inaddition, there are thermodynamic limitations in conventionaldehydrogenation. These thermodynamic limitations are due to the factthat conversion in conventional dehydrogenation processes areequilibrium limited, and require high temperature and low pressure toachieve high single pass yields. As a result of these substantialdrawbacks, the petroleum industry has sought a solution to the demandfor olefinic hydrocarbons in the use of autothermal cracking andoxidative dehydrogenation methods.

In autothermal cracking, oxygen or air is added to the feed andpartially combusts part of the feed in situ generating the hightemperatures required to thermally crack the remaining feedstock. Insome variants a catalyst is used to support combustion with the catalystbeing in the form of a fixed bed or a fluidized or spouted bed. Fixedbeds are preferred to reduce catalyst attrition. In some cases hydrogenis co-fed with the feedstock and is found to increase olefin yields.Autothermal cracking usually takes place at high temperatures (550-1200°C.) and requires very short reaction times and rapid quenching of theproducts to preserve the olefinic products and prevent furtherundesirable reactions. Even so, by products are formed including carbonoxides. At higher pressure, yields of undesirable by-products increase.At very high temperatures as encountered in some autothermal processes,hydrocarbon cracking to methane also reduces selectivity to usefulolefinic products.

Catalytic oxidative dehydrogenation is, in principle, not subject tomany of the problems associated with conventional steam cracking orcatalytic dehydrogenation because of the presence of oxygen in thereaction mixture. Oxidative dehydrogenation (ODH) uses oxygen to reactwith the hydrogen released from the hydrocarbon, in situ, so that theaforementioned equilibrium limitation is removed, and high single passyields can be achieved. The reaction is exothermic overall and does notrequire a supply of heat as in endothermic dehydrogenation reactions.Generally, in a catalytic oxidative dehydrogenation process, thereactants (hydrocarbon and an oxygen-containing gas) are passed over thefixed bed catalyst directly to produce olefin product. Typically, thehydrocarbon is a saturated hydrocarbon such as ethane or a mixture ofsaturated hydrocarbons. The hydrocarbon may be gaseous or liquid atambient temperature and pressure but is typically gaseous.

An example of an alkene which can be formed via an oxidativedehydrogenation process, is ethylene. The latter process is attractivefor many reasons. For example, compared to thermal cracking, high ethaneconversion can be achieved at moderate temperatures (300-1000° C.) bycatalytic oxidative dehydrogenation. Unlike thermal cracking andcatalytic dehydrogenation, catalytic ODH is exothermic, requiring noadditional heat, beyond feed pre-heat, to sustain reaction. Furthermore,in contrast to catalytic dehydrogenation, catalyst deactivation by cokeformation should be minimal in ODH because of the presence of oxygen inthe reactor feed. Other alkanes can similarly be oxidativelydehydrogenated.

Although there are no reported commercial ODH processes operating at thepresent time, there is a high level of commercial interest. Activity hasfocused on ethane, propane and isobutane ODH, and patents to same haveissued. Representative of these patents are the following US patents,all of which are herein incorporated by reference: U.S. Pat. Nos.4,524,236; 5,162,578; 5,593,935; 5,997,826; 6,313,063; 6,281,378;6,239,325; 6,235,678; 6,130,183; 6,355,854 and 6,310,241.

Industrial interest has stimulated investigations into new catalysts andmethods for improved performance (e.g., conversion and selectivity) forthe oxidative dehydrogenation of alkanes. U.S. Pat. No. 4,524,236reports high conversion (73%) and high selectivity (71%) for ethane ODHbut these results were obtained only by diluting the ethane/oxygen feedwith helium as 85.5% of the feedstock. Others have achieved high yieldsby co-feeding hydrogen with the hydrocarbon feedstock and oxygen (seeU.S. Pat. No. 5,997,826).

In U.S. Pat. No. 4,524,236 McCain describes a process for the lowtemperature catalytic oxydehydrogenation of ethane to ethylene in a gasphase and featuring the use of a catalyst containing Mo/V/Nb/Sb and anadditional element.

There have been different approaches to adding oxygen to the ODHreaction. Lodeng et al. in U.S. Pat. No. 5,997,826 describes a processfor converting C3 and C4 paraffins to olefins by a sequential reactorthat contains at least three zones, a catalytic dehydrogenation processzone, an oxygen admixing zone, and a catalytic oxidation zone, whereinthe flow velocity in the admixing zone is higher than in the catalystzones. Ward in U.S. Pat. No. 4,739,124 discloses mixing oxygen betweenstages.

In a process for catalytic selective oxidation of a hydrocarbon,Perregaard et al. in U.S. Pat. No. 6,515,146 discloses a reactor inwhich oxygen flows into a 7 mm inner diameter tube through the porousalumina tube walls and into the catalyst bed held within the tubes. Nomention is made of the useful of this approach in ODH.

Beretta et al. in “Production of olefins via oxidative dehydrogenationof light paraffins at short contact times,” Catalysis Today, 64 pp103-111 (2001) reported testing of a Pt/Al₂O₃/Fe—Cr catalyst in anannular reactor. Comparative tests without catalyst showed no proof thatthe Pt catalyst contributed to the selective oxidation of ethane toethene; however, there was strong proof “that the catalyst was active innon-selective oxidation reactions, and that gas-phase oxidativepyrolysis was a fast process with very high ethene selectivities.” Theauthors concluded that the Pt-containing catalyst seemed to be mainlyactive in the total oxidation of ethane to CO_(x).

Several workers have described oxidative dehydrogenation in catalystmonoliths positioned in conventional reactors. See, U.S. Pat. Nos.4,940,826, 6,166,283, and 6,365,543. They do not suggest the use ofmonoliths in microchannel reactors or any microchannel advantages.

As compared to conventional, fixed bed reactors, microchannel reactorshave been found to suppress thermal gradients; however, at comparablecatalyst bed temperatures, the microchannel reactor did not improveperformance. Steinfeldt et al. in “Comparative Studies of the OxidativeDehydrogenation of Propane in Micro-Channels Reactor Module andFixed-Bed Reactor,” Studies in Surface Science and Catalysis, pp 185-190(2001) conducted testing of ODH in a microchannel reactor over aVOx/Al₂O₃ catalyst. To minimize temperature gradients, the catalyst wasdiluted with quartz in a ratio of 1:9. The authors reported that “theuse of micro-channels reactor module allowed isothermal operation at allreaction conditions.” The authors concluded that the “micro-channelreactor module and fixed bed reactor show approximately the samecatalytic results under isothermal conditions.”

Despite extensive research, there remains a need for new oxidativedehydrogenation catalysts, catalytic systems, and methods that achievehigh conversion at high selectivity, such that the yield of the desiredolefin is maximized, and extraneous oxidative side reactions areminimized. Such extraneous oxidative side reactions may include theconversion of starting hydrocarbon, e.g., alkane, into carbon oxides (COand/or CO₂), and/or conversion of desired product alkene into carbonoxides.

SUMMARY OF THE INVENTION

According to the scientific literature, the performance of the oxidativedehydrogenation reaction in microchannels did not differ from thereaction in conventional fixed bed reactors operating at the sametemperature. Thus, in view of the fact that microchannel apparatus isgenerally more expensive than conventional equipment, there appeared tobe no reason to conduct oxidative dehydrogenation in microchannelapparatus. Despite this discouraging background, we proceeded to testoxidative dehydrogenation reactions in microchannel reactors.Surprisingly, we found that conducting the oxidative dehydrogenationreaction in microchannel apparatus produced significantly superiorresults as compared to the same reaction in larger, moreconventionally-sized apparatus.

In one aspect, the invention provides a method for catalytic oxidativedehydrogenation of a gaseous hydrocarbon, comprising: flowing ahydrocarbon-containing fluid and a source of oxygen into a microchannel;wherein a catalyst is present in the microchannel; reacting thehydrocarbon-containing fluid and the source of oxygen, in themicrochannel, in a temperature range of 335 to 1000° C., to form waterand at least one alkene and/or aralkene; and removing heat into anadjacent heat exchanger. In preferred embodiments, heat is removed intoan adjacent heat exchanger by (a) cooling the microchannel by flowing acoolant fluid (which could be a reactant stream) through an adjacentcooling chamber to convectively cool the reaction microchannel, or (b)conducting a simultaneous endothermic reaction in adjacent channel(s) toremove heat, or (c) performing a phase change in adjacent channel(s),preferably microchannel(s), to provide additional heat removal beyondthat provided by convective heat exchange in adjacent channel(s).

In another aspect, the invention provides a method for catalyticoxidative dehydrogenation of a gaseous hydrocarbon, comprising: flowinga hydrocarbon-containing fluid and a source of oxygen into amicrochannel; wherein a catalyst is present in the microchannel;reacting the hydrocarbon-containing fluid and the source of oxygen, inthe microchannel, in a temperature range of 335 to 1000° C., to formwater and at least one alkene and/or aralkene; and quenching the streamformed after reacting the hydrocarbon-containing fluid and the source ofoxygen.

In a further aspect, the invention provides a method for catalyticoxidative dehydrogenation of a gaseous hydrocarbon, comprising: flowinga hydrocarbon-containing fluid and a source of oxygen into amicrochannel; wherein a catalyst is present in the microchannel;reacting the hydrocarbon-containing fluid and the source of oxygen, inthe microchannel, in a temperature range of 335 to 1000° C., to formwater and at least one alkene and/or aralkene; and feeding oxygen intothe microchannel at multiple points along the channel length.

In another aspect, the invention provides a method for catalyticoxidative dehydrogenation of a gaseous hydrocarbon, comprising: flowinga hydrocarbon-containing fluid and a source of oxygen into amicrochannel; wherein a catalyst is present in the microchannel;reacting the hydrocarbon-containing fluid and the source of oxygen, inthe microchannel, in a temperature range of 335 to 1000° C., to formwater and at least one alkene and/or aralkene; and wherein said methodis characterized by superior conversion, selectivity and/or yield, suchthat, as compared to a reaction conducted under the same conditions(reactant feed composition, oxidant, diluent, ratios offeed/oxidant/diluent (with diluent level as close as practicable),contact time, pressure, catalyst bed temperature, catalyst compositionand form) in a 1.0 cm inner diameter quartz tube with no active coolingand pre-mixed hydrocarbon and oxidant (that is, no staged oxidant), theresults of the method exhibits one or more of the following: (a) an atleast 20% relative higher ratio of selectivities of CO/CO₂; or (b) an atleast 10% relative higher conversion of hydrocarbon; or (c) an at least10% relative higher yield of olefins; or (d) an at least 10% relativehigher selectivity to olefins; or (e) an at least 10% relative lowerselectivity of carbon dioxide. By “relative” is meant in comparison tothe quartz tube, for example, if the method in a quartz tube produced a10% conversion, an 11% conversion would be 10% higher relativeconversion. This method differs from ODH through a monolith in aconventional reactor that would not necessarily produce enhancedresults. In contrast, persons ordinarily skilled in this technologywould, in light of the teachings set forth herein, be able through nomore than routine experimentation to identify suitable operatingconditions to obtain the claimed enhanced results.

In another aspect, the invention provides a method of oxidativelydehydrogenating a gaseous hydrocarbon, comprising: flowing ahydrocarbon-containing fluid and a source of oxygen into a microchannel;wherein an oxidative dehydrogenation catalyst is present in themicrochannel; reacting the hydrocarbon-containing fluid and the sourceof oxygen, in the microchannel, in a temperature range of 300 to 1000°C., to form water and at least one alkene and/or aralkene; wherein thehydrocarbon comprises an alkane or aralkane, and wherein diluent, ifpresent, constitutes 0.25 or less, as a volume fraction, of total fluidflow through the microchannel.

In a further aspect, the invention provides a method of oxidativelydehydrogenating a gaseous hydrocarbon, comprising: flowing ahydrocarbon-containing fluid and a source of oxygen into a microchannel;wherein an oxidative dehydrogenation catalyst is present in themicrochannel; reacting the hydrocarbon-containing fluid and the sourceof oxygen, in the microchannel, in a temperature range of 300 to 1000°C., to form water and at least one alkene and/or aralkene; wherein thehydrocarbon comprises an alkane or aralkane, and wherein at least 10% ofthe hydrocarbon is converted to an alkene and/or aralkene; and whereintotal hydrocarbon feed flow through the microchannel is at a LHSV ofabout 32 or greater.

In a further aspect, the invention provides a method of oxidativelydehydrogenating a gaseous hydrocarbon, comprising: flowing ahydrocarbon-containing fluid and a source of oxygen into a microchannel;wherein an oxidative dehydrogenation catalyst is present in themicrochannel; reacting the hydrocarbon-containing fluid and the sourceof oxygen, in the microchannel, in a temperature range of 300 to 1000°C., to form water and at least one alkene and/or aralkene; wherein thehydrocarbon comprises an alkane or aralkane, and wherein at least 10% ofthe hydrocarbon is converted to an alkene and/or aralkene; and whereindiluent, if present, constitutes 0.25 or less, as a volume fraction, oftotal fluid flow through the microchannel, and wherein total hydrocarbonfeed flow through the microchannel is at a LHSV of about 1 or greater.

The invention also provides a method of oxidatively dehydrogenating agaseous hydrocarbon with reduced gas phase reactions, comprising:flowing a hydrocarbon-containing fluid and a source of oxygen into amicrochannel; wherein an oxidative dehydrogenation catalyst is presentin the microchannel; and wherein the hydrocarbon-containing fluid andthe source of oxygen are combined immediately before contacting thecatalyst such that precatalyst contact time is 150 ms or less.

In yet another aspect, the invention provides a method of oxidativelydehydrogenating a gaseous hydrocarbon with reduced gas phase reactions,comprising: flowing a hydrocarbon-containing fluid and a source ofoxygen into a microchannel; wherein an oxidative dehydrogenationcatalyst is present in the microchannel; and wherein the combinedpressure of hydrocarbon-containing fluid and the source of oxygen in afeed stream is at least 10 atmospheres (when measured under standardconditions) and the precatalyst contact time of thehydrocarbon-containing fluid or the source of oxygen at a temperature of300° C. or more is 15 ms or less.

In another aspect, the invention provides a method of oxidativelydehydrogenating a gaseous hydrocarbon, comprising: flowing ahydrocarbon-containing fluid and a source of oxygen into a reactionchamber; wherein an oxidative dehydrogenation catalyst is present in thereaction chamber; wherein the oxidative dehydrogenation catalystcomprises an oxide catalyst comprising Mg, V and Mo, wherein the molarratio of Mo:V is in the range of 0.5 to 2; reacting thehydrocarbon-containing fluid and the source of oxygen, in the reactionchamber, to form water and at least one alkene and/or aralkene.

In another aspect, the invention provides apparatus for oxidativelydehydrogenating a hydrocarbon, comprising: a microchannel reactionchamber; and an oxidative dehydrogenation catalyst disposed in themicrochannel reaction chamber; and comprising: an oxygen channeladjacent to said microchannel reaction chamber and separated by anoxygen channel wall, wherein apertures through said oxygen channel wallform passageways between the oxygen channel and the reaction chamber.

In another aspect, the invention provides apparatus for oxidativelydehydrogenating a hydrocarbon, comprising: a microchannel reactionchamber; and an oxidative dehydrogenation catalyst disposed in themicrochannel reaction chamber comprises one of the following forms:

-   -   a) a particulate catalyst; or    -   b) a porous insert; or    -   c) a catalyst wall coating comprising a first layer formed        between a reaction chamber wall and a second layer; wherein the        reaction chamber wall, first layer and second layer have        different compositions, wherein the first layer has a thickness        of at least 0.1 micrometers, more preferably at least 1.0        micrometers.

In a further aspect, the invention provides a catalytic system foroxidatively dehydrogenating a hydrocarbon, comprising: a reactionchamber; and an oxidative dehydrogenation catalyst disposed in thereaction chamber; wherein the system is characterizable by a catalyticactivity such that when propane and O₂, with no diluents, in a 1:1 ratioare fed into the reaction chamber at an LHSV of 32 and a catalysttemperature of 580° C., there is a propane conversion of at least 30%and an olefin yield of at least 20%.

As exemplified in the aspect above, any of the systems and methods can,in some cases, be characterized in conjunction with properties such asconversion, yield and/or selectivity. These properties can be selectedfrom any of the values in the descriptions of preferred embodiments orfrom the data in the Examples section.

Advantages provided by various embodiments of the present invention mayinclude one or more of the following: relatively high levels ofalkane(s) and/or aralkane(s) conversion and high selectivity toalkene(s) and/or aralkene(s); relatively low selectivity to by-products,such as carbon monoxide or carbon dioxide; and the ability to conductoxidative dehydrogenation without diluents added to either the feedand/or the catalyst—thus providing a more efficient and compacttechnique.

Other advantages of the process of the present invention include:maximization of intercontact of the source of oxygen, the hydrocarbon,and the catalyst material; and, minimization of homogenous gas-phaseunselective reactions, such as those which convert starting and/orproduct hydrocarbon to carbon oxides (CO_(x)).

Further advantages which may accrue to the processes of the presentinvention include the possibility of process intensification.Conventional ODH and autothermal cracking processes of the prior art areoften operated under condition of reactant dilution to prevent runawayreactions (and prevent explosions), while the process of the presentinvention can be operated, if desired, under more intensive conditionsleading to greater throughput. By combining catalytic microchannel andadjacent heat exchangers it is possible to operate at feed/oxygen ratiosthat would conventionally lead to high temperatures and loss ofselectivity, but by removing heat rapidly through heat exchange with theheat removal channels, the temperature in the catalytic channels can bekept relatively low (in some embodiments below 700° C., or below 600°C., or below 500° C.), thus maximizing selectivity to desired olefinproducts. The inventive process can be operated very nearly in athermally neutral mode, wherein the heat released by the oxidationchemistry very nearly matches the heat consumed in the crackingreactions, thus minimizing the need to remove or add large amounts ofheat to the reactor.

GLOSSARY

“Adjacent” means directly adjacent such that a wall separates twochannels or chambers; this wall may vary in thickness; however,“adjacent” chambers are not separated by an intervening chamber thatwould interfere with heat transfer between the chambers.

By “including” is meant “comprising”, however, it will be understoodthat the terms “consists of” or “consists essentially of”, mayalternatively be used in place of “comprising” or “including” todescribe more limited aspects of the invention.

“Integrated” means all the components are within the same structurewherein the exhaust zones are directly connected to the reactionchambers.

Liquid hourly space velocity (LHSV) is defined based on the liquidvolumetric flow and the reaction chamber volume. Reaction chamber volumeis defined as the volume of a process channel where catalyst is presentand the temperature is sufficiently high for dehydrogenation to occur.Reaction chamber volume is the wall-to-wall volume and includes catalystvolume (including pore volume, and, if present, interstitial volume),and, if present, the volume of a bulk flow path or paths through or bythe catalyst. For dehydrogenation of isobutene, a “sufficiently high”temperature will typically be at least about 400° C., fordehydrogenation of propane, typically at least about 450° C. Tocalculate LHSV, GHSV (h-1), defined as volumetric flow rate of gas ofhydrocarbon (ml/h) per volume catalyst (ml), is calculated and then itis divided by a factor that relates the volume of a quantity of the feedin the gas phase to the volume of the same quantity of the feed as aliquid (230 for propane). This factor takes into account the differencein the density of the hydrocarbon in liquid and gas phase.

Contact time is calculated as 3600/GHSV (hydrocarbon) and has dimensionsof seconds. Contact time is defined by the volume of the reactionchamber divided by the volumetric feed flow rate of the reactantcomposition. The volumetric feed flow rate is the sum of the hydrocarboninlet flow rate(s) and the inlet oxidant flow rate(s) taken as if theywere gasses at a temperature of 0° C. and a pressure of 1 atmosphere.

“A reactant stream containing a hydrocarbon” can also be termed “ahydrocarbon stream,” and, in the context of the present invention, theseterms mean the entire gas stream (not merely a selected portion thereof)entering a reaction chamber(s).

“ODH” is oxidative dehydrogenation.

“Autothermal cracking” is oxidative dehydrogenation which requiresminimal or no net heat input to or removal from the system.

“Thermally neutral” means a process in which the difference between theenthalpy of the product mixture leaving the reactor zone and theenthalpies of the reactants entering the reactor zone (including oxidantand diluent) is less than 25% (in some embodiments 10% or less and insome embodiments 5% or less) of the combined reactant enthalpies. Insome embodiments, the methods described herein are thermally neutral.

Definitions of the performance parameters used herein, are as follows.“Percent conversion” refers to the moles of carbon in the organiccompound(s) to be dehydrogenated (e.g., moles of carbon in the alkane)that are consumed, based on the moles of carbon in the said organiccompound(s) fed to the reactor. “Percent selectivity” refers to themoles of carbon in the products (e.g., alkene) formed based on the molesof carbon consumed. “Percent yield” refers to the moles of carbon in thedesired product(s) (e.g., alkene) formed based on the moles of carbonfed. For reaction mixtures of ethane, propane or butane, desiredproducts are ethene, propene, and butenes, respectively. Percentselectivity and percent yield are based on carbon. To give ahypothetical example, a reaction mixture containing 2 moles of hexaneand 1 mole ethane that results in a product mixture containing 1 molehexane, 1 mole ethene, 0.5 mole hexene, 2 mole CO₂ and 0.33 mole propenewould have a 57% carbon conversion with a (6 mol C)/(8 mol C)=75%selectivity to olefins (37.5% hexene, 12.5% propene, 25% ethene) and42.8% yield of olefins (21.4% hexene yield, 7.1% propene, 14.3% ethene).

BRIEF DESCRIPTION OF THE DRAWINGS

FIG. 1A illustrates a cross-flow microchannel reactor for ODH. ODHcatalysts can be placed in the process channels as either a coating, aninsertable felt or foam, or packed powders. A coolant couldalternatively be oriented as co-flow or counter-flow. The coolant couldbe a gas stream, a liquid stream such as hot oil or molten salt, a phasechange liquid, or an endothermic reaction such as reforming.

FIG. 1B is a schematic illustration of flow through a reaction chamber.

FIGS. 2A and 2B are schematic illustrations of integrated reactordesigns showing the process and heat exchange channels and flows.

FIGS. 3A-3C are schematic illustrations of integrated reactor designsshowing the process and heat exchange channels with distributed flow.

FIGS. 4A and 4B are schematic illustrations of integrated reactordesigns with recuperative heat exchange between process streams.

FIGS. 5A and 5B are schematic illustrations of integrated reactordesigns that are “numbered up” to achieve greater capacity.

FIGS. 6 and 7 illustrate catalysis testing devices described in theExamples section.

DESCRIPTION OF THE INVENTION

The invention includes a method for catalytic oxidative dehydrogenationof a hydrocarbon. In this method, a hydrocarbon-containing mixture (themixture is or contains a fluid and may be homogeneous or heterogeneous(for example, containing some colloidal liquid droplets or solidparticulates)) flows past and/or through a catalyst material. Preferablythe mixture is entirely gaseous. The mixture comprises a source ofoxygen and at least one hydrocarbon; in preferred embodiments, theoxygen source is introduced immediately before the catalyst zone orwithin the reactor catalyst zone or in a staged fashion. A portion ofthe at least one hydrocarbon reacts to form at least one alkene and/oraralkene and a portion of the source of oxygen reacts to form water.Optionally, the product stream can be rapidly quenched to preserveproducts and stop further reaction to undesirable products. Quenchingcan be achieved using integral microchannel quench/heat exchanger toremove heat in adjacent channels to the channels through which theproduct is flowing. Heat exchange can be between the product stream andthe feed stream. In another embodiment, the quench can be achieved bymixing the hot product stream with a cold fluid to rapidly reducetemperature. The quench fluid can be condensible fluids, for example,excess low temperature steam or a condensible hydrocarbon injected as aliquid that evaporates and cools the product stream by absorbing latentheat of evaporation from the hot product stream. Condensible fluids areattractive for use in commercial applications for gas-phase products,since they are relatively easily separated from the product mixture.

This invention discloses methods for the oxidative dehydrogenation ofalkane(s) and/or aralkane(s) to alkene(s), alkadiene(s) and/oraralkene(s). The hydrocarbon may be any alkane or aralkane of C₂ up toC₂₀. Examples of alkane include ethane, propane, isobutane or butane orhigher alkanes including up to C₂₀ linear and branched alkanes; examplesof aralkane include ethylbenzene; examples of alkene for the purpose ofthis invention include ethylene, propylene and also alkadienes such asbutadiene; examples of aralkene include styrene. Preferred examples ofhydrocarbons are C₂-C₁₈ alkanes, preferably C₂-C₁₀ alkanes, isobutane,propane, ethane, ethylbenzene, or C₁₀-C₁₅ alkanes such as could be usedfor making detergent alcohols. The alkanes can be linear, branched andcyclic. Hydrocarbons can be obtained commercially either in pure form orin mixtures. Hydrocarbons can also be derived from other reactions, andthe output of these reactions used with or without an interveningpurification step. Systems of the invention can be described asincluding apparatus and/or catalyst in combination with reactants and/orproducts. By “including” is meant “comprising”, however, it will beunderstood that any of the terms “consists of” or “consists essentiallyof”, may alternatively be used to describe more limited aspects of theinvention. Additionally, any of the individual components (such asethane, for example) may preferably be present in at least 20% purity(based on carbon atoms), or at least 50%, or at least 90%, or 100%purity.

The source of oxygen is preferably a gas capable of providing molecularoxygen, which may conveniently be molecular oxygen or air. Oxygen (O₂)is preferred over air, and in preferred embodiments, the O₂:N₂ ratio (orthe O₂:diluent ratio) entering a reaction chamber is one or greater,more preferably at least 3, and still more preferably at least 10. Insome embodiments, the hydrocarbon/oxygen (O₂) ratio in the feedpreferably is 2.0 or more, in some embodiments between 1 and 3, in someembodiments 1.8 or less, in some embodiments 2.5 or more.

For autothermal ODH of ethane to ethylene the ethane:H₂ feed ratio ispreferably in the range 1:0 to 1:1, preferably 1:0.2 to 1:0.6, mostpreferably 1:0.25 to 1:0.5, and the ethane:O2 feed ratio should remainin the range 1:0.1 to 1:1, preferably 1:0.2 to 1:0.8 and most preferably1:0.25 to 1:0.5 depending on the overall reaction selectivities andconversion.

The reactant stream may contain diluents such as nitrogen, methane,water vapor, CO, and CO₂. Steam, if present in the reactant feed, ispreferably present in a steam:C ratio of 5 or less, more preferably 1 orless, and in some embodiments 2 volume % or less. The total diluents todehydrogenatable hydrocarbons molar ratio is preferably 5:1 or less,more preferably 2:1 or less, preferably less than 50 volume %, morepreferably less than 20 volume % diluents in a microchannel reactor, andin some embodiments, less than 2 vol. % diluents. In some preferredembodiments, the hydrocarbons in the reactant stream are at least 75 mol%, more preferably at least 90 mol % of a single hydrocarbon (propane,for example). In some preferred embodiments, the reaction contains nodiluent except H₂. In order to enhance selectivity, optional hydrogenmay be co-fed with the starting hydrocarbon. The hydrogen may be fedfrom a separate source or produced in the ODH reaction and recycled. Insome embodiments, there is no H₂ in the reactant stream, in someembodiments there is a 0 to 5H₂:hydrocarbon ratio on a molar basis.

Microchannel reactors are characterized by the presence of at least onereaction channel having a (wall-to-wall, not counting catalyst)dimension of 2.0 mm (preferably 1.0 mm) or less, and in some embodiments50 to 500 μm. Both height and width are perpendicular to the directionof flow. The height and/or width of the reaction microchannel ispreferably 2 mm or less, and more preferably 1 mm or less (in which casethe reaction chamber falls within the classical definition of amicrochannel). The length of the reaction channel is typically longer.Preferably, the length of the reaction chamber is greater than 1 cm,more preferably in the range of 1 to 25 cm. Typically, the sides of thereaction channel are defined by reaction channel walls. These walls arepreferably made of a hard material such as a ceramic, an iron basedalloy such as steel, or monel. More preferably, the reaction chamberwalls are comprised of stainless steel or inconel which is durable andhas good thermal conductivity. The reactors can be made by knownmethods, and in some preferred embodiments are made by laminatinginterleaved shims, where shims designed for reaction channels areinterleaved with shims designed for heat exchange. A “shim” is a thinflat sheet that optionally has voids to create flow channels or paths.

The reactors preferably include a plurality of microchannel reactionchannels and/or a plurality of adjacent heat exchange microchannels. Theplurality of microchannel reaction channels may contain, for example, 2,10, 100, 1000 or more channels. In preferred embodiments, themicrochannels are arranged in parallel arrays of planar microchannels.During operation, the heat exchange microchannels contain flowingheating and/or cooling fluids. Non-limiting examples of this type ofknown reactor usable in the present invention include those of themicrocomponent sheet architecture variety (for example, a laminate withmicrochannels) exemplified in U.S. Pat. Nos. 6,200,536 and 6,219,973(both of which are hereby incorporated by reference). Performanceadvantages in the use of this type of reactor architecture for thepurposes of the present invention include their relatively large heatand mass transfer rates, and the ability to safely operate in explosiveregimes. Unlike conventional reaction vessels for ODH, which have totake into account the possibility of explosion for mixtures of oxygenand hydrocarbon, this is advantageously less of a possibility in theprocess of the present invention. Furthermore, use of microchannelreactors can achieve better temperature control, and maintain arelatively more isothermal profile, compared to architectures of theprior art. This, in turn, advantageously leads to lessened peaktemperatures and lessened coking of the hydrocarbon starting materialand/or desired product. Lower peak temperatures also reduce unselectivehomogeneous gas phase reations leading to carbon oxides.

An example of microchannel reactor hardware suitable for ODH is shown inFIG. 1A. Coolant microchannels (typically 2 mm or less) are adjacent toa microchannel reaction chamber (2 mm or less). The wall between thechannels is preferably 2 mm or less. The flow of coolant may be orientedin a co-current flow, counter-current flow, or cross-current flow. Thelength of the process flow channel may be any length, but a typicalrange is 1 to about 10 inches (2.5 to 25 cm). The height of the processchannel may also be any value, but a typical range is 0.1 inches toabout 10 inches (0.25 to 25 cm). Each of the process or coolant channelmay be further subdivided with parallel subchannels. The spacing ofsubchannels is dependent upon maximizing heat transfer and minimizingmechanical stresses.

An alternate microchannel design for ODH reactions is the close couplingof an endothermic reaction in an adjacent microchannel. The placement ofan endothermic reaction such as a steam reforming reaction next to theexothermic ODH reaction allows for the highest rate of heat transfer. Atypical heat flux for convective cooling in a microchannel reactor is onthe order of 1-5 W/cm2. The incorporation of a simultaneous endothermicreaction to provide an improved heat sink will enable a typical heatflux of roughly an order of magnitude above the convective cooling heatflux.

A simplified representational view of an apparatus of some embodimentsof the present invention is illustrated in FIG. 1B. The views shown inthe figures are representative examples and should not be understood tolimit the invention. A process channel 2 contains a bulk flow path 4.The reaction chamber is defined on two sides by reaction chamber walls 6and 6′. The internal dimension h (height) is the distance from thesurface of the metal wall 8 to the surface of the metal in the opposingwall and does not include the thickness of any oxide layer (not shown).A heating chamber 10 is adjacent to process channel 2. The illustratedheating chamber has fins 11 having a thickness d interleaved withheating channels 14 and a gap 12 between the fins and the channel wall6. In preferred embodiments, the distance between fins and/or thethickness of the heating chamber is 2 mm, more preferably 1 mm or less.The illustrated embodiment is cross-flow; however, co-flow andcounter-flow may also be employed. In some preferred embodiments, anendothermic reaction is occurring in the cooling channel; however, acool, non-reacting stream could alternatively be used. In someembodiments, the heating chamber 10 is divided into several parts, forexample regions 7, 9, 13 into which various fluids could flow to tailorthe temperature profile in a process channel. For example, steam or thereturn portion of a ODH stream could flow through region 7 to provide apreheat zone; an endothermic process stream can flow through region 9 toremove heat from the oxidative dehydrogenation reaction in a reactionchamber (a portion of the process channel in which catalyst 15 ispresent), and a cold fluid flows through region 13 to quench thereaction.

Another schematic illustration of a cross-section of an integratedreactor design is illustrated in FIG. 2A. A reactant stream(hydrocarbon) and oxygen source flows into the inlet (Fluid B inlet) ofa forward process channel, passes through a u-turn, and then flows inthe reverse direction in the return process channel. At the same time, aheat transfer fluid flows into the inlet (Fluid A inlet) of a heattransfer channel, passes through a u-turn, and then flows in the reversedirection in the return heat transfer channel. It is desirable to matchthe coolest portion of the heat transfer channel with the reactionchamber portion of the process channel. In a preferred embodiment, thereaction chamber is located in the return process channel in an area 23located near the u-turn (closer to the u-turn than the outlet) so thatthe reactant stream flowing through the forward process channel 25 iswarmed by the return process stream (which could be termed the “exhaust”(i.e., the product stream) and the reaction chamber). More preferably,the heat transfer fluid is an endothermic reaction stream that reacts ina catalyst-containing portion located in the return heat transferchannel in an area 27 located near the u-turn opposite the ODH reactionchamber; in which case the endothermic reaction stream in the forwardheat transfer channel 29 is preheated by the ODH chamber (the area wherethere is endothermic reaction catalyst and an endothermic reactionoccurs) and exhaust stream. This type of reactor design is especiallydesirable where the u-turn end 24 (i.e., the hot end) is relativelyunconstricted so that it can expand when the device is in operation,manifolds can be connected at the inlet end (i.e., the cold end). As istrue of all the reactor designs described herein, the illustratedreactor can be stacked to increase reactor capacity; for example threeof the illustrated reactors can be stacked in a single integrated deviceto have six layers: heat exchange: process: heat exchange: process: heatexchange: process; preferably with all the inlets and outlets located onone side of the device. In some preferred embodiments, the u-turnsconnect to a single return channel and are not manifolded.

An alternative design, particularly advantageous for operation when thefeed gas is adjusted so that the reaction is nearly thermally neutral,excludes the heat exchange channels, increasing reactor capacity. Inthis embodiment the channels are not interlayered with heat exchangechannels but are preferably arranged as stacks of hydrocarbon feed,oxygen feed and product channels, repeating this pattern multiple times.

An alternative design is illustrated in FIG. 2B in which return channels26, 28 are disposed between forward channels. The operation of thisdevice is analogous with the reactor of FIG. 2A, except in preferredembodiments the respective catalysts are located in the forward process30, 31 and heat exchange channels 32, 33 near the u-turns. Although thecatalysts are depicted as partially filling a cross-section of a processchannel (such catalysts could be, for example, catalytic inserts), ODHcatalysts could fill a cross-section of a process channel (such as, forexample, a packed bed). ODH catalysts preferably comprise a wallcoating.

Reactor designs illustrating the distributed flow concept areillustrated in FIGS. 3A-3C. In distributed flow, a secondary fluidenters into a reaction chamber. FIG. 3A illustrates a device in which afirst fluid (Fluid B) flows through a first channel 35. Adjacent to thischannel is a second channel 36 into which feeds Fluid A. Fluid C entersthe reactor in a separate channel 37 and then flows in a distributedfashion through apertures 38 along the length of the second channel. Insome embodiments, the first channel contains a oxidative dehydrogenationcatalyst (not shown) and a hydrocarbon and oxidant flows into thechannel. In some embodiments, the second channel contains an endothermiccatalyst (not shown) and either a hydrocarbon or an oxidant flows intothe inlet of the second channel (Fluid A Inlet) while, at the same time,another reactant flows into a third channel (Fluid C Inlet) and flowsthrough apertures 38 into the endothermic reaction chamber where anendothermic reaction occurs, in some embodiments there is an endothermiccatalyst on wall 301 and the endothermic reaction occurs at or near thewall separating the first and second channels. This controls the rate ofendothermic reaction and matches the heat generation rate with the heatrequired to drive the endothermic reaction. Any thermal profile can betailored.

Alternatively, a heat transfer fluid (Fluid B) can pass through thefirst channel. In some preferred embodiments, the first channel 35contains an endothermic catalyst (not shown) and Fluid B contains anendothermic mixture. A reactant (hydrocarbon) can flow in through eitherinlet (Fluid A Inlet or Fluid C Inlet) and react over a (oxidativedehydrogenation) catalyst in the second channel 36. When hydrocarbon(optionally containing an oxidant) enters into the third channel 37(through Fluid C Inlet) it flows in a distributed fashion into thesecond channel for a controlled reaction over the length of the reactionchamber; in this case, a secondary fluid flows through the secondchannel. Alternatively, a (hydrocarbon) reactant stream enters throughFluid A Inlet while an oxidant enters Fluid C Inlet and flows into thereaction chamber in a distributed fashion through the apertures. Thesecondary fluid can be reactive (such as an oxidant) or a nonreactivediluent. A nonreactive diluent can quench the reaction. A cold secondaryfluid can be effective in rapidly quenching a reaction.

Alternative designs are illustrated in FIGS. 3B and 3C in which flowscan be controlled as have been described in FIG. 2 and FIG. 3A. Channelshave been illustrated as open channels but it should be recognized thatthe channels may contain features such as catalysts, microchannelgrooves, and/or support ribs. The illustrated designs assume the typicalsituation in which the oxidative dehydrogenation reaction is exothermic;however, the invention also includes embodiments in which the reactionis heat-balanced, that is, neither exothermic or endothermic butsufficient oxidation occurs to supply just enough heat to drive thedehydrogenation reactions and make up for heat loss to the environment.In this case the heat exchange channels can optionally be eliminated. Insome embodiments, oxidative dehydrogenation occurs to an extentsufficient to have a significant effect on the overall reaction but theoxidation reactions don't generate sufficient heat to make up for thelosses to the dehydrogenation reactions and to the environment, in thiscase, heat needs to be added to the ODH process channel and this heatcould be provided, for example, from a hot fluid (or exothermicreaction) in an adjacent heat exchange channel.

Another way to integrate heat exchange in an integrated reactor isillustrated schematically in FIGS. 4A and 4B. In this embodiment, afirst reactant stream (Fluid A1, containing a hydrocarbon and oxidant)flows in a first direction (dashed arrow 47) through a first processchannel 41 while a second reactant stream (Fluid A2, containing ahydrocarbon and oxidant) flows in an opposite direction (dashed arrow46) in a second process channel. Heat exchange is provided to bothprocess channels via an intervening, cross-flow heat exchange channel43. Preferably, an ODH catalyst 44, 45 is disposed within each processchannel 41, 42 on the process channel wall that is adjacent the heatexchange channel to form a reaction chamber within each process channel.Catalyst can optionally be coated on any or all of the walls of theprocess channel. In cases where a reaction occurs in the heat exchangechannels a catalyst may optionally be placed in these channels. The hotproduct stream exiting the reaction chamber is immediately quenched bythermal transfer with the incoming reactant stream in the adjacentprocess channel. The illustrated embodiments show the process channelsas separated by a constant distance; however, it should be appreciatedthat the process channels could be positioned closer to each other inthe recuperation zones (i.e., the zones where the process channels areadjacent, that is, the zones without an intervening heat exchangechannel). Assigning length as the direction parallel to flow within eachchannel and height as the one direction that is perpendicular to flow inboth the process channels and the heat exchange channel, and width beingthe remaining dimension, it is preferred that the length of each processchannel be at least three times, more preferably 10 times longer thanthe width of the heat exchange channel; and, preferably, the preheatzone of the first process channel is of substantially the same length asthe quench or “exhaust” zone of the second process channel, and viceversa. Preferably, the length of the preheat zone of each processchamber is preferably at least as long as the width of the heat exchangechannel; similarly, the length of the quench zone of each processchamber is preferably at least as long as the width of the heat exchangechannel. It can readily be appreciated that the capacity of this type ofdevice can be increased by stacking up to any desired height withalternating heat exchange and process channels; in some embodiments atleast 3 of each.

Sheets of channels and/or integrated reactors can be “numbered up” toobtain greater capacity. A schematic illustration of an exploded view ofa stack of three identical sheets is shown in FIG. 5A. In a deviceformed by laminating these three sheets, a first fluid (such as a heatedfluid) flows into inlet 53 through the first and third sheets and exitsvia outlet 55 while a process stream 57 (for example, containing ahydrocarbon) flows through the second sheet. In this figure, the darkregions indicate a solid material, while the white areas indicate areasfor fluid flow (such as could be formed by etching). Flow occurs throughall the channels. To further increase capacity, blocks 51 of multi-levelreactors (see FIG. 5B) can be manifolded and operated together.

It is advantageous to reduce temperature of the product stream asrapidly as possible after leaving the catalyst section of themicrochannel reactor to prevent further undesirable reactions of theolefins. This rapid cooling is known as “quenching.” An integrated orseparate heat exchanger can be used to quench the reaction products,cooling them down rapidly once the reaction has taken place. Forexample, near the outlet of a reaction channel, cross-flow coolantchannels can rapidly cool the product stream. In some preferredembodiments, the heat from the product stream is transferred to areactant stream in a microchannel heat exchanger, thus preheating ahydrocarbon stream that can be subsequently dehydrogenated. The heatfrom the product stream could also be used to drive an endothermicreaction. Another form of quench is the rapid addition of a reactive(such as reactant feed) or a non-reactive gas into the hot productstream; this could be accomplished through a gas inlet or inlets locatedin a reaction chamber, or in or near a reaction chamber outlet, and,optionally with the aid of a static mixer structure within thedownstream pipe.

In several of the methods and reaction systems described herein, thereaction products are quickly quenched. Thus, the reaction zone may beclosely and integrally linked with a heat exchange zone (eitherrecuperative or other) to quickly cool the reaction mixture after thereactor to below 300° C. or by rapid mixing with secondary, cooler gasstream. Integrated microchannel heat exchanger(s) preferably cool thereaction mixture at a rate greater than 1° C. per millisecond of averageheat exchanger residence time; more preferably, at a rate greater than5° C. per millisecond of average heat exchanger residence time. In somepreferred embodiments, the temperature of the process stream decreasesby 100, more preferably 200 and still more preferably 300° C. within 50milliseconds (ms), more preferably 10 ms after reacting (that is, afterpassing through the hot reaction zone), and in some embodiments 1 ms to500 ms, preferably 1 ms to 100 ms. Temperatures in reactionmicrochannels can be measured with thermocouples.

In some embodiments of the inventive reactor or method, the reactor (ormethod) is configured to send the product stream into a second reactoror recycle the product stream back into the same reactor. There may beintervening separation steps to remove desired products or undesiredcomponents or separate hydrogen or a reactant or reactants. In somepreferred embodiments, separation is conducted within the sameintegrated device as the dehydrogenation. Typically, the desired alkeneor arylalkene will be separated from the product stream and theunreacted hydrocarbon portion of the product stream recycled.

A product stream containing olefins and unconverted alkanes can be usedwithout further separation as a feedstock for other processes includingalkylation. In alkylation, (typically) olefins are reacted withisoalkanes to form higher branched alkanes with high octane numberssuitable for use as components of gasoline. Where the feedstock containsisobutane, the product stream is especially suited as an alkylationfeedstock since the products include C3-C5 olefins and unconvertedisobutane.

In some preferred embodiments, walls of the reaction channels and/orinner surfaces of conduits and manifolds connected to the reactionchannels are coated with a passivation layer. Passivation of surfacesinside the reaction chamber and/or in piping leading to, and/orespecially piping leading from the reaction chamber may reduce cokingand nonselective oxidation reactions and might enhance time-on-streamperformance. Passivation coatings have a different composition than theunderlying material. Suitable passivation coatings include a refractoryoxide such as silica, alumina, zirconia, titania, chromia, ceria, GroupII metals (alkaline earths) and rare earth metals, atomic numbers 57-71.It has been unexpectedly discovered that a silica coating demonstratedsuperior selectivity compared to an alumina passivation layer. Thepassivation coating could, optionally, be catalytic supports or could bedense coatings to protect an underlying metal wall. Passivation coatingscan be made by applying a sol, or a fine particulate coating onto ametal surface, or applied by chemical or physical vapor deposition orelectrochemical deposition, or thermally-grown, or combinations of thesetechniques. It is believed that surfaces quench undesired gas phaseunselective oxidations. Thus, insome embodiments, filler material 17such as ceramic fibers are placed into the reaction channel in openspaces within the reaction channel that, during operation, would beoccupied by hot gas. The filler quenches gas phase reactions and thusimproves selectivity.

The reaction channel contains an oxidative dehydrogenation catalyst.Suitable catalyst structures within the reaction channel include porouscatalyst materials, monoliths, washcoats, pellets, and powders. Thecatalyst can comprise a high surface area support and an overlying layeror layers comprising a catalytically active metal or metals. In somepreferred embodiments, the reaction is cooled by an adjacent endothermicreaction stream and, in some embodiments, an adjacent heat exchangechannel comprises a catalyst that may contain structures such as porouscatalyst materials, monoliths, washcoats, pellets, and powders.

The catalytically-active material of the process of the presentinvention is not particularly limited and may include any effectiveprior art ODH catalyst. Among the catalytically-active materials of thepresent invention are the so-called high-temperature catalysts, i.e.,those comprising noble metals, preferably catalyst materials comprisingat least one metal selected from the group consisting of Pt, Pd, Rh, Irand Ru. Also among the catalytically-active materials of the presentinvention are the so-called low-temperature catalysts, which maycomprise at least one oxide or phosphate of a metal selected from thegroup consisting of Li, Mo, V, Nb, Sb, Sn, Zr, Cr, Mg, Mn, Ni, Co, Ce,rare-earth metals (such as Sm), and mixtures thereof. The low or hightemperature catalyst may contain additional components such as alkalaior alkaline earth promoters, or metals such as Cu, Ag, or Sn. Preferredsupport materials include alumina, silica, other metal oxides,mesoporous materials and refractory materials. Examples of some suitablecatalyst compositions are described in U.S. Pat. Nos. 6,130,183 and5,997,826. Catalysts can be, for example, vanadia dispersed on alumina,or platinum on alumina. Catalysts can also be a noble metal dispersed ona metal oxide layer that is coated over (such as by wash coating orchemical vapor deposition) a metal foam or metal felt (nonwoven metal).In some preferred embodiments, catalyst is disposed (such as by CVD orwash coating) on a wall or walls of a microchannel.

The catalyst can fill up a cross-section of the reaction channel (aflow-through catalyst) or only occupy a portion of the cross-section ofa reaction channel (flow-by). In a flow-by catalyst configuration, gaspreferably flows in a 0.1-1.0 mm gap adjacent to a porous insert or athin layer of catalyst that contacts the microchannel wall (in someembodiments, the microchannel wall that contacts the catalyst is indirect thermal contact with a heat exchanger, typically, in theseembodiments, a heated fluid or exothermic reaction process streamcontacts the opposite side of the wall that contacts the catalyst).

In embodiments, the reaction channel contains a porous catalyst materialthat defines at least a portion of at least one wall of a bulk flowpath. In this preferred embodiment, the surface of the catalyst definesat least one wall of a bulk flow path through which the mixture passes.During operation, the mixture flows through the microchannel, past andin contact with the catalyst. The term “bulk flow path” refers to anopen path (contiguous bulk flow region) within the reaction chamber. Acontiguous bulk flow region allows rapid gas flow through the reactionchamber without large pressure drops. In preferred embodiments there islaminar flow in the bulk flow region. Bulk flow regions within eachreaction channel preferably have a cross-sectional area of 5×10⁻⁸ to1×10⁻² m², more preferably 5×10⁻⁷ to 1×10⁻⁴ m², and the maximum distancefrom the mid-stream of the bulk flow path is less than 1 mm, preferablyless than 0.6 mm. The bulk flow regions preferably comprise at least 5%,more preferably 30-99% of either 1) the internal volume of the reactionchamber, or 2) the cross-section of the reaction channel.

In some preferred embodiments, the catalyst is provided as a porousinsert that can be inserted into (or removed from) each channel in asingle piece; preferably the porous insert is sized to fit within amicrochannel with a width of less than 2 mm. In some embodiments, theporous catalyst occupies at least 60%, in some embodiments at least 90%,of a cross-sectional area of a microchannel. In another embodiment, thecatalyst is a coating (such as a washcoat) of material within amicrochannel reaction channel or channels.

A “porous catalyst material” (or “porous catalyst”) refers to a porousmaterial having a pore volume of 5 to 98%, more preferably 30 to 95% ofthe total porous material's volume. At least 20% (more preferably atleast 50%) of the material's pore volume is composed of pores in thesize (diameter) range of 0.1 to 300 microns, more preferably 0.3 to 200microns, and still more preferably 1 to 100 microns. Pore volume andpore size distribution are measured by Mercury porisimetry (assumingcylindrical geometry of the pores) and nitrogen adsorption. As is known,mercury porisimetry and nitrogen adsorption are complementary techniqueswith mercury porisimetry being more accurate for measuring large poresizes (larger than 30 nm) and nitrogen adsorption more accurate forsmall pores (less than 50 nm). Pore sizes in the range of about 0.1 to300 microns enable molecules to diffuse molecularly through thematerials under most gas phase catalysis conditions. The porous materialcan itself be a catalyst, but more preferably the porous materialcomprises a metal, ceramic or composite support having a layer or layersof a catalyst material or materials deposited thereon. The porosity canbe geometrically regular as in a honeycomb or parallel pore structure,or porosity may be geometrically tortuous or random. Preferably thesupport is a foam metal or foam ceramic. The catalyst layers, ifpresent, are preferably also porous. The average pore size (volumeaverage) of the catalyst layer(s) is preferably smaller than the averagepore size of the support. The average pore sizes in the catalystlayer(s) disposed upon the support preferably ranges from 10⁻⁹ m to 10⁻⁷m as measured by N₂ adsorption with BET method. More preferably, atleast 50 volume % of the total pore volume is composed of pores in thesize range of 10⁻⁹ m to 10⁻⁷ m in diameter. Diffusion within these smallpores in the catalyst layer(s) is typically Knudsen in nature, wherebythe molecules collide with the walls of the pores more frequently thanwith other gas phase molecules.

At a point where the chamber height or the chamber width is about 2 mmor less, the chamber height and the chamber width define across-sectional area. In some preferred embodiments, the cross-sectionalarea comprises a porous catalyst material and an open area, where theporous catalyst material occupies 5% to 95% of the cross-sectional areaand where the open area occupies 5% to 95% of the cross-sectional area.In some preferred embodiments, the open area in the cross-sectional areaoccupies a contiguous area of 5×10⁻⁸ to 1×10⁻² m². In other preferredembodiments, the catalyst occupies greater than 98% of thecross-sectional area.

If necessary, the catalyst systems can be regenerated by treating thecatalyst with an oxidant to oxidize reduced materials formed on or inthe catalyst. Typical regeneration oxidants are oxygen or air. Catalystscan be refurbished after irreversible reduction of activity byimpregnating or coating the catalyst in situ with additional activematerials.

In addition to the reaction microchannel(s), additional features such asmicrochannel or non-microchannel heat exchangers may be present.Microchannel heat exchangers are preferred. An integrated or separateheat exchanger can be used to quench the reaction products, cooling themdown rapidly once the reaction has taken place to prevent furtherundesirable reactions of the olefins. In some embodiments of theinventive reactor or method, the reactor (or method) is configured tosend the product stream into a second reactor or recycle the productstream back into the same reactor. Adjacent heat transfer microchannelsenable temperature in the reaction channel to be controlled to promoteselective ODH and minimize unselective reactions in the gas phase thatincrease with temperature. The heat exchange fluids can be gases orliquids and may include steam, liquid metals, or any other known heatexchange fluids—the system can be optimized to have a phase change inthe heat exchanger. In some preferred embodiments, multiple heatexchange layers are interleaved with multiple reaction microchannels(for example, at least 10 heat exchanger layers interleaved with atleast 10 reaction microchannel layers, where heat exchanger layers areseparated by about 2 mm or less.

Many other options exist for the design of a microchannel reactor. Forexample, a process channel can be in thermal contact with a productchannel, an oxygen feed channel, or both. In a simple embodiment, acoolant gas flows in adjacent microchannels to the ODH reaction chamber.The flow of coolant may be cross flow, counter-flow or co-flow. Coflowmay be preferred to obtain the greatest heat flux in the beginning of areaction chamber if the process reaction will be greatest at the frontof the reaction chamber where reactants are most concentrated. Animprovement to heat transfer would be the use of a higher heat capacityfluid, such as a molten salt or a hot oil. The hot oil coolant istypically limited to systems with reaction temperatures no greater than400° C. and the molten salts would be used for much higher temperatures.

In an alternate microchannel embodiment, the air or oxygen used for theODH reaction could be staged or fed sequentially into the reactionmixture. The staging could occur in separate devices, through the use ofsmall orifices or jets within one device, or from a microporous membraneor alternate sparging sheet. The controlled addition of oxygen topartial oxidation reactions, and specifically oxidative dehydrogenationreactions, has been demonstrated in the literature (Tonkovich, Zilka,Jimenz, Roberts, and Cox, 1996, “Experimental Investigations ofInorganic Membrane Reactors: a Distributed Feed Approach for PartialOxidation Reactions”, Chemical Engineering Science, 51(5), 789-806) fordistributed feed membrane reactors. Staged oxygen addition (i.e.,distributed oxygen feed) lowers the local oxygen partial pressure andthus favors the desired partial oxidation reaction over the competingand undesired combustion reaction. Literature experimental and modelingresults demonstrate this effect for series-parallel reactions. Thestaged addition may also limit peak temperatures by leveling oxygenconcentration through the reaction zone.

In some preferred embodiments, an adjacent channel (or channels) carriesan oxygen source that is distributed over a length of the reactionmicrochannel(s). In some preferred embodiments, oxidant enters areaction chamber at more than 3 points along the chamber length. In someembodiments where a reaction chamber is defined by plural walls(typically four), there are oxidant inlets on one, or more than one,wall of the reaction chamber. The inlets need not be uniformlydistributed along the length of the reaction chamber, but positioned foroptimal results.

In some preferred embodiments, the hydrocarbon(s) and oxygen-source aremixed such as by a microchannel mixer that is separate or integral withthe reaction microchannel. Mixing is preferably conducted beforereaction but can be conducted during reaction such as by a mixerdisposed within a reaction microchannel.

There is no particular limit on pressure of the reaction. For bettereconomy, pressure should be relatively high. In some preferredembodiments, pressure of the feed is at least 50 kPa, more preferably atleast 100 kPa. In some preferred embodiments, pressure of reactants(i.e., excluding partial pressure of diluents) is greater than 1 atm,more preferably greater than 2 atm. Pressure of the feed should bemeasured prior to contact with the ODH catalyst. In some embodiments,pressure in the reactor is 10 bar or less. In some embodiments, pressuredrop through the reactor, or through a reaction channel, is 2 bar orless, an in some embodiments 0.5 bar or less.

Hydrocarbon oxydehydrogenation is conducted at modest pressure, about 1atm or less, in conventional reactors. Attempts to increase the pressureto higher pressures, greater than 1 atm, 2 atm or more, result in sharpreductions in selectivity to the desired alkene or arylalkene products.As the pressure of the reactant mixture is increased the intensity ofheat release, that is the heat release per unit volume, increasesproportionally, and the rates of various oxidation reactions increasewith the increased partial pressures of the reacting gases. Thus aspressure increases in a conventional reactor the local heat releaseincreases and, due to the limited capability of the conventional reactorto remove heat, the temperature rises. Thus it is not possible tooperate conventional reactors at high pressures and high space velocity.With microchannel reactors the high heat removal capacity makes itpossible to run reactions at higher pressures and high space velocityand still achieve high selectivity at high conversion. With pressuresabove 2 atm, preferably above 5 atm, and more preferably above 10 atmand space velocities greater than 10,000 h-1, preferably greater than100,000 h-1, and more preferably greater than 1,000,000 h-1 it ispossible to get good yields of useful products in microchannel reactors.

Hydrocarbon to oxygen ratios in oxidative dehydrogenation reactions aresubject to limitations for various reasons. Mixtures containing oxygenand hydrocarbons can be explosive. Indeed, consideration of explosivelimits are an important facet of safe plant and process design.Explosive limits become narrower, i.e., more limiting in terms of theacceptable oxygen to hydrocarbon ratio, as pressure increases. Thenarrower limit of the explosive regime at higher pressure can preventsafe operation of processes at high pressure. Microchannel reactorsprovide the opportunity to operate in regimes that might otherwise beconsidered unsafe due to explosive limit considerations. In themicrochannel reactor only very small volume mixtures of oxygen andhydrocarbon are available within any one connected region, for exampleone channel, so that explosions are not expected to propagate amongseparate channels. Furthermore, the dimensions of the microchannelreactors are similar to the so-called quench diameters of manyoxygen/hydrocarbon mixtures. At dimensions below the quench diameter theradical chain reactions that cause explosions are terminated by contactwith the device wall, eliminating explosions or flames. Flame arrestorswork on this principle. In ODH in microchannels, it is possible to workat pressures above 1 atm, preferably above 2 atm, more preferably above5 atm, most preferably above 10 atm with high oxygen to hydrocarbonvolume ratios, even within the explosive or flammable regimes, oxygen tohydrocarbon ratios greater than 0.2:1, greater than 0.3:1, greater than0.5:1, even greater than 1:1, without diluents and without explosivereactions.

Preferred temperature ranges of the process of the present inventioninclude: a temperature ranging from 335 to 1000° C., more preferably500-900 C, and in some embodiments about 500 to about 700. Preferably,during operation, temperature of the catalyst and the adjacent reactionchamber wall differ by less than 10° C.

Gas hourly space velocity (GHSV) of the inventive methods preferablyrange from 1,000 h⁻¹ to 10,000,000 h⁻¹ based on reactor volume, or 1,000ml feed/(g catalyst) (hr) to 10,000,000 ml feed/(g catalyst) (hr). Inother preferred embodiments, GHSV is at least 10,000 h⁻¹ or at least10,000 ml feed/(g catalyst) (hr); more preferably at least 100,000 h⁻¹or at least 100,000 ml feed/(g catalyst) (hr); more preferably at least500,000 h⁻¹ or at least 500,000 ml feed/g catalyst; more preferably atleast 1,000,000 h⁻¹ or at least 1,000,000 ml feed/(g catalyst) (hr).

Liquid hourly space velocity (LHSV) is preferably at least as fast asthe examples, e.g., at least 4 h⁻¹; more preferably at least 16 h⁻¹;more preferably at least 64 h⁻¹; more preferably at least 127 h⁻¹.Contact times in the reaction chamber (the catalyst zone) preferably arein the range of 0.001 to 5 s, more preferably less than 500 ms, morepreferably less than 100 ms, and still more preferably less than about70 ms. As shown in the following Examples section, we surprisingly foundthan as contact time decreases, the total olefin yield increases forpropane ODH conducted in a microchannel even in the face of stable ordecreasing propane selectivity. When the C3:O2 ratio was 2:1 and theprocess inlet temperature was 540° C. and catalyst bed temperature was538° C. total olefin yield increased from 21% at 1470 ms contact time to30.6% at 367 ms contact time. When propane ODH was conducted in amicrochannel at a C3:O2 ratio of 2:1 and the process inlet temperaturewas 597° C. and catalyst bed temperature was approximately 600° C.,total olefin yield increases from 29.3% at 250 ms contact time to 33.1%at 82 ms contact time and to 37.8% at 61 ms.

In preferred embodiments employing a quench step, the sum contact timesin the precatalyst zone, the catalyst zone and the quench zone ispreferably 1 second or less, more preferably 500 ms or less, morepreferably 200 ms, and still more preferably 100 ms or less.

The amount of heat that can be transferred through a plane separatingthe process reaction chamber from a heat exchanger is a function of themethod of heat transfer. For convective heat transfer from a hot fluidin a heat exchange channel to a dehydrogenation reaction chamber, theamount of heat (as defined as Watts per square can of reaction chamberwall area that is adjacent to the heat exchanger) transferred for agaseous heat transfer fluid is preferably at least 1 W/cm² and may be upto about 15 W/cm² For a liquid heat transfer fluid used in convectiveheat transfer, higher heat transfer fluxes are achievable and may rangefrom at least 1 W/cm² to about 30 W/cm². For conductive heat transferfrom an exothermic reaction, much higher rates of heat transfer areattainable and heat flux may range from about 10 W/cm² to about 100W/cm². These defined ranges of heat fluxes are for steady-stateoperation and average over the area of a process reaction chamber wallthat is adjacent to a heat exchanger; or, in a reactor with multiplechannels (more than two channels), an average over the areas of alldehydrogenation reaction chambers adjacent to heat exchanger(s) in allthe channels in operation.

Preferably, selectivity to carbon oxides (on a carbon atom basis) isless than 40%, more preferably less than 20% (in some embodiments, inthe range of 20% and 5%), and even more preferably less than 5%. In lesspreferred embodiments, selectivity to carbon dioxide (on a carbon atombasis) is less than 40%, more preferably less than 20% (in someembodiments, in the range of 20% and 5%), and even more preferably lessthan 5%.

The CO/CO₂ ratio is indicative of the efficiency of the ODH process; lowratios indicate that oxygen was unavailable for ODH and was consumedprimarily for combustion. In a microchannel reactor we are capable ofobtaining CO to CO₂ ratios in excess of those predicted at equilibriumfor the particular gas mixture in question when the reactor temperatureis below the temperature at which the formation of CO is favoured overthe formation of CO₂. For example when the ratio of C3 to O2 is 2:1 andthe total pressure is 10 psig the temperature at which CO and CO₂ are ata 1:1 ratio at equilibrium is approximately 660° C. below thistemperature the formation of CO is strongly favored by thermodynamics.

For a given mixture at a given operating pressure the CO:CO2 ratioobtained in a microchannel reactor when the temperature is below thatwhere the formation of CO is thermodynamically favorable and ispreferably at least 2.4:1 or more preferably 2.76:1 or more preferably4.6:1 or even more preferably 10:1.

At equal peak temperatures the volumetric productivity as defined by thegrams of target olefin (for example propylene) produced per unit volumeof reaction chamber (reaction chamber is that portion of a channel wherecatalyst is present either as flow-by or flow-through) per hour isgreater in a microchannel than in a conventional reactor. As shown inthe examples, when the C3 to O2 ratio was 1:1 and the peak temperaturewas about 625° C. the productivity of the microchannel is greater thanthat of a quartz tube by a factor of 1.9. Volumetric productivity of amicrochannel reactor performing propane ODH could in one instance be 15g/ml/hr or preferably 30 g/ml/hr or more preferably 60 g/ml/hr or evenmore preferably 120 g/ml/hr or more, in some embodiments productivity is15 to about 150 g/ml/hr.

In the case of ethane ODH, at equal average temperatures, theproductivity as defined by the grams of target olefin (for exampleethylene) produced per unit mass of catalyst of catalyst per hour isgreater in the microchannel than in a conventional reactor. When the C2to O2 ratio was 10:1, the oxidant was air and the average temperaturewas close to 650° C. the productivity of the microchannel was found tobe greater than that of a quartz tube by a factor of 7.4. Productivityof a microchannel reactor performing ethane ODH is preferably at least270 g/g/hr or more preferably at least 600 g/g/hr or more preferably1200 g/g/hr or even more preferably at least 2400 g/g/hr.

Once oxidant has been mixed with the hydrocarbon the potential existsfor unwanted oxidations to occur (i.e. the production of CO and CO₂).The injection of the oxidant into the hydrocarbon stream (or vice versa)just upstream of the catalyst (as was done in the ODH v2 microchannelpellet—see Examples) has the potential to reduce these reactionsespecially once the gasses are at or near reaction temperature (˜400°C.). A pre-catalyst contact time based on the volume between the firstpoint at which the oxidant contacts the hydrocarbon and the point atwhich the catalyst starts is preferably less than 150 ms or preferablyless than 75 ms or more preferably less than 40 ms or even morepreferably less than 10 ms.

The rates at which the undesirable combustion reactions proceed aredependent on the total pressure with increased pressure leading toincreased rate of reaction in addition they are dependent on the oxidantpartial pressure that also increases if the C3:O2 ratio is fixed and thetotal pressure is increased. These undesirable reactions can beminimized and the selectivities to the desirable olefins maintained ifthe total inlet pressure is at least 10 atm and the contact time in thepre-catalyst zone is less than 15 ms or preferably less than 7.5 ms ormore preferably less than 4.0 ms or even more preferably less than 1 ms.Undesirable reactions can be minimized and the selectivities to thedesirable olefins maintained if in a another embodiment the total inletpressure is at least 20 atm and the contact time in the pre-catalystzone is less than 7.5 ms or preferably less than 4.0 ms or morepreferably less than 2.0 ms or even more preferably less than 0.5 ms. Inanother embodiment, undesirable reactions can be minimized and theselectivities to the desirable olefins maintained if the total inletpressure is at least 30 atm and the contact time in the pre-catalystzone is less than 4.0 ms or preferably less than 2.0 ms or morepreferably less than 1.0 ms or even more preferably less than 0.25 ms.

The percent conversion of hydrocarbon (in a single pass) is preferably10% or higher, more preferably about 20% or higher, more preferably 40%or higher, even more preferably 50% or higher. The level of percentselectivity to desired product or products in the case where more thanone valuable alkene can be formed, is preferably at least 10% preferablyat least 20%, preferably at least 40%, and in some embodiments 10 toabout 60%. The yield of product alkene or alkenes and/or aralkene in mol% per cycle is preferably greater than 10%, and more preferably greaterthan 20%. The total yield of product alkene or alkenes and/oraralkene(s), in mol %, is preferably greater than 50%, more preferablygreater than 75%, and most preferably greater than 85%. The specifiedlevels of conversion, yield and selectivity should be understood asexemplary and include all values such as yield per cycle of at least15%, at least 25%, etc. as well as ranges such as 10 to 30%. The rangesand conditions can be further understood with reference to the Examplesand the invention includes all ranges and minimum levels of conversions,etc. described therein. It is also envisioned that routine testing andexperimentation, in view of the teachings provided herein, will revealsuperior results and it is therefore intended that this disclosure bebroadly interpreted to include descriptions of numerous levels (andranges) of conditions and results.

Oxygen conversions of greater than 90%, greater than 95%, mostpreferably greater than 99% can be achieved with gas flow rates ofgreater than 10,000 h-1, greater than 100,000 h-1 and even greater than1,000,000 h-1 in an oxidative dehydrogenation process in a microchannelreactor.

While preferred embodiments of the present invention have beendescribed, it will be apparent to those skilled in the art that manychanges and modifications may be made without departing from theinvention in its broader aspects. The appended claims are thereforeintended to cover all such changes and modifications as fall within thetrue spirit and scope of the invention.

EXAMPLES Description of Devices

Throughout the Examples section, the term “pellet” does not have itsusual meaning, but takes a special meaning of a microchannel testingapparatus as described here. The “ODH v1” microchannel test pellets weredesigned to provide active cooling on both sides of the processmicrochannel and use pre-mixed feeds. The ODH v1 pellets were fabricatedfrom 2.8″ (7.1 cm) long piece of 0.75″ (1.9 cm) Inconel™ 625 bar stockusing a combination of wire EDM, plunge EDM, conventional machining andwelding. Each device contained 3 microchannels, 1 process microchannelsandwiched between 2 cooling microchannels. Each ran axially and thechannels were parallel in alternating planes 0.040″ (0.10 mm) apart. Theprocess channel in each device had the dimensions 0.020″×0.300″×2.65″(0.050×0.762×6.73 cm). Each cooling channel had the dimensions0.020″×0.400″×2.038″ (0.050×1.02×5.17 cm). The cooling microchannelswere formed by first opening the channel for the entire length of thedevice, 2.8″ (7.1 cm), and were then isolated from the process channelin the header and footer region by the insertion of 0.395″×0.306″×0.020″plugs that were subsequently seam welded in place (thus reducing flowlength of the channels from 2.65″ to 2.038″. Access to the coolingchannels was obtained at the inlet and outlet ends of the pellet bymachining a 0.43″ diameter blind hole normal to the major axis of thepellet the depth of which, 0.175″, being sufficient to break into thecoolant channels but not the reactant. Both header and footer werelocated on the same face. Washtub type headers were welded in place overthe holes at each end. Tubing to allow for the attachment of fittingswas welded to the inlet and outlet faces of the tube body and to thecoolant headers and footers. An illustration of the device is providedin FIG. 6. Thermowells for 0.020″ thermocouples were provided atlocations 0.400″, 1.067″, 1.733″ and 2.400″ from the inlet face of theprocess microchannel on the face opposite of the oxidant header andfooter. The ODH v.1 pellets were run with heat treatment and passivationlayers and without heat treatment or passivation layers.

The ODH v.2 pellet was designed to allow active cooling on one face ofthe microchannel as well as to have the capacity to introduce theoxidant into the reactant in the process microchannel via opposed jetsjust upstream from the catalyst resulting in a well mixed reactantstream entering the reaction zone. The body of the pellet was fabricatedfrom a 0.625″×0.500″×2.700″ piece of Inconel™ 617 using a combination ofwire EDM, plunge EDM, conventional machining and welding (see FIG. 7). Amicrochannel with an opening of 0.025″×0.300″ was then cut into thebody. The center line of the channel being 0.250″ down the 0.625″ sideof the 0.500″×0.625″ face thus the process microchannel is off-centertoward the ‘top’ of the pellet. The coolant channel is formed by cuttinga 0.313″ deep by 0.300″ wide by 1.500″ long pocket in the bottom of thepellet 0.225″ from one end of the pellet. The end of the pellet closestto the pocket is then defined as the outlet side of the pellet. Thecoolant channel was formed when a second piece with a 0.283″ tenon isinserted leaving a channel that was 0.030″×0.300″ by 1.500″. Thecoolant/process web is 0.0495″ thick. Inlet and outlet ports for thecoolant are provided at each end of the tenon piece via 0.069″ throughholes.

Material was removed from the top of the pellet to a depth of 0.187″starting at 0.287″ from the inlet face and ending 0.288″ from the outletface. Material was removed from the bottom face to a depth of 0.313″starting at 0.287″ from the inlet face and ending 0.787″ from the inletface. The opposing oxygen jets are formed by putting 5 0.020″ diameterthrough holes in the top of the piece. The first two holes have theircentres located on a line 0.400″ downstream of the inlet face and 0.077″to each side of the device's axial centre line. The center of the thirdhole is located 0.533″ downstream from the inlet face on the axialcenter line of the device. The final 2 holes have their centres locatedon a line 0.667″ downstream of the inlet face and 0.077″ to each side ofthe device's axial center line. The oxidant header was formed from awashtub type header, 0.490″×0.644″×0.875″, with a 0.134″×0.500″ slot.This piece was fitted over the area containing the jets and welded inplace. Welding on a 0.490″×0.644″×0.063″ plate sealed the open face.

Thermowells for 0.020″ thermocouples were provided at locations 0.288″,0.725″, 1.225″ and 1.662″ from the outlet face of the processmicrochannel on the face opposite of the oxidant header and footer.

Prior to use the ODHv.2 devices were cleaned by sonication for 20 min inhexane bath followed by immersion in 20% HNO₃ solution for 20 minutes.After the cleaning step the device was subjected to heat treatmentprotocol (see Table 1).

After heat treatment, the surfaces of the device that come into contactwith the reactant hydrocarbons were passivated with an alumina sol-coat(Dispal 14N4-25). This was done by forcing the alumina sol through thereactant inlet tubing, through the reactant microchannel and into theproduct footer and allowing the sol to remain in contact with thesurfaces for 15 min. Excess sol was then removed using nitrogen purgestream flowing through the reactant inlet and, to ensure the jets didnot become blocked, simultaneously through the oxidant inlet. Aftercoating the device was calcined by heating to 200° C. at a rate of 1°C./min (to allow it to dry slowly) and then heated to 1000° C. and heldthere for 1 hour.

TABLE 1 Heat Treatment Protocol employed for ODH v2 pellets StepTemperature Ramp Rate Flow/Atmosphere Total Time Comments 1 Ambient 0200 SCCM/N₂ as needed Vacuum chamber & device 3x and replenish N₂ 2 900° C. 3.5° C./min 84 SCCM/H₂ & 200 SCCM/  4 hr N₂ flows via bubblerwith heat N₂ with H₂O 15 min tape @ 60° C. vapour 3  900° C. hold ~1SLPM/N₂ 30 min 4 1000° C. 3.5° C./min 200 SCCM/N₂ 29 min 5 1000° C. Hold200 SCCM/Air  1 hr Air @ 1 SLPM for first 10 min 6  25° C. 3.5° C./min200 SCCM/Air  4 hr 38 min

Two versions of the ‘ODH v3’ microchannel test pellets were fabricatedfrom 0.5″ diameter Inconel 617 rod via a combination plunge and wire EDMand conventional machining. In ODH v3a the process channel was formed bycutting a 0.020″×0.375″ axial slot in a 2.003″ long piece of the rodmaterial. In ODH v3b the process channel was formed by cutting a0.035″×0.370″ axial slot in a 2.003″ long piece of the rod material. Inboth versions, a 0.73″ long by 0.19″ deep pocket was then cut on eitherside of the process channel thus leaving 0.0425″ between the inner wallsof the process channel and the outside of the pellet.

ODH v3a/b pellets were heat-treated following the protocol in Table 2.After heat-treating a sol-coat of alumina was applied to the surface.This was done by forcing the alumina sol (Dispal 14N4-25) into theprocess channel and allowing the sol to remain in contact with thesurfaces for 15 min. After coating the devices were calcined by heatingto 200° C. at a rate of 1° C./min and then heating to 1000° C. with ahold time at the peak temperature of 1 hour. The ODH v3a pellet wastested first with an alumina sol coat and subsequently with a silica solcoat. To form the silica coat, the alumina coat on the ODH v3a pelletwas removed by sonication in saturated solution of NaOH for severalhours, followed by cleaning the metal surface with acetone. The bareInconel 617 pellet was then dip-coated with a silica sol (hydrolysedtetraethyl orthosilicate (TEOS) with HNO₃) and left to gel for 48 h inair at room temperature. The pellet was then further dried at 80° C. for1 h followed by calcination (formation of glass) at 1000° C. for 3 h.

TABLE 2 Heat Treatment Protocol employed for ODH v3 Pellets StepTemperature Ramp Rate Flow/Atmosphere Total Time Comments 1 Ambient 0200 SCCM/N₂ as needed Vacuum chamber & device 3x and replenish N₂ 2 900° C. 3.5° C./min 84 SCCM/H₂ & 200 SCCM/  4 hr N₂ flows via bubblerwith heat N₂ with H₂O 15 min tape @ 60° C. vapor 3  900° C. hold ~ 1SLPM/N₂ 30 min 4 1000° C. 3.5° C./min 200 SCCM/N₂ 29 min 5 1000° C. Hold200 SCCM/Air 10 hr Air @ 1 SLPM for first 10 min 6  25° C. 3.5° C./min200 SCCM/Air  4 hr 38 min

Inconel 625 tubes (0.25″ nominal O.D., I.D. 0.188″) were employed in theethane ODH testing. Prior to use these tubes were heated to 1000° C. instagnant air and held at this temperature for 4 hours. Scale was removedvia sonication.

Example 1 Propane Oxidative Dehydrogenation in Conventional andMicrochannel Reactors

The ODH catalyst used in all the testing of this example was an Mg—V—Ocatalyst containing 80.9 wt % MgO and 19.1 wt % V₂O₅ with surface areaof 98 m²/g. Catalyst was pelleted to a size between 250-400 μm (pressedto 5 tons, ground and sieved off the desired fraction) and thenpre-treated prior to reaction at 500° C. in 40 ml/min O₂ for 1 h. InTable 3, “quartz” refers to a 1 cm inner diameter quartz tube containingthe packed catalyst.

The exothermicity of the reaction was followed by a thermocouple placedat the bottom of the catalyst bed. For safety reasons, at the beginningof the reaction, nitrogen was introduced into the feed mimicking ODH inair (O₂:N₂=1:4). Later, in a stepwise way, the diluent was pulled out(ratios of 1:3, 1:2 and 1:1) until it was completely removed. At eachstep, GC analysis of the reaction effluent was done after 5 mintime-on-stream.

TABLE 3 Effect of LHSV on Conversion and Selectivity in the aMicrochannel Pellet C₃:O₂ = 1:1, ODH v1 Blank Quartz Device Quartz TubeODH v1* Quartz Tube ODH v1* ODH v1* ODH v1* tube LHSV (v/v//hr) 4 4 3232  62 157 62 C3 Conversion (%) 49.2 55.0 38.7 48.0  43.3 79.9 57.6Propylene Yield (%) 21.6 16.9 15.0 11.5  9.7 13.0 13.5 Olefin Yield (%)30.9 29.0 18.6 22.1  18.5 43.6 31.0 CO_(x) Selectivity (%) 30.0 36.850.7 42.7  45.3 22.1 27.6 CO/CO₂ Ratio 1.02 1.19 1.57 1.11  1.37 4.556.90 O₂ Conversion (%) 66.8 67.3 72.3 66.4  61.2 36.9 39.9 Selectivityto CO₂(%) 14.8 16.8 19.7 20.2  19.1 4.0 6.1 Selectivity to CO (%) 15.120.0 31.0 22.5  26.2 18.2 21.4 Selectivity to CH₄(%) 6.6 9.5 1.1 10.0 10.9 19.7 16.9 Selectivity to C₂H₆(%) 0.6 0.9 0.1 1.2  1.0 3.6 1.7Selectivity to Propylene (%) 43.9 30.8 38.9 23.9  22.5 16.3 23.4Selectivity to Ethylene (%) 18.9 21.9 9.2 22.2  20.3 38.3 30.5 CatalystBed Temp (° C.) 540 540 636 540 542** 512 538 *was not heat treated orgiven surface coating prior to operationThe results in Table 3 above at a C3:O2 ratio of 1 and an inlet gastemperature of 540° C., show that at the lowest LHSV=4 conversion in themicrochannel device is higher than the fixed bed and that overall olefinyield is approximately the same. At LHSV=32, the temperature in theconventional reactor rises to 636° C. while that in the microchannelremains at 540° C. due to the better heat removal properties of themicrochannel. Despite the lower temperature in the microchannel,conversion in the microchannel is higher and olefin yield is alsohigher. At LHSV=157 in the microchannel device there is a slighttemperature drop to 512° C., while the conversion and olefin yieldincrease substantially. Normally at higher throughput (higher LHSV) itwould be expected that temperature would increase but in themicrochannel the temperature falls slightly. It would also be expectedthat conversion would fall off at higher LHSVs and not increase as itdoes here. This surprising result indicates that the microchannel deviceis operating in a different mode to the conventional fixed bed andproducing unexpectedly high yields of useful olefins.

A series of reactions were conducted at ratios of O₂:N₂ of 4:1. 3:1,2:1, 1:1, and 1:0 at about 545° C. at varying space velocities in the“blank” quartz tube and the catalyst-containing microchannel reactor. Inthe blank tube, hydrocarbon conversion remained at less than 10% withpropene selectivity of about 60% until diluent was removed (O₂:N₂=1:0),at which point conversion jumped to 45% while propene selectivity fellto 23% and ethylene selectivity rose from about 7% to about 28%. In thecatalyst-containing microchannel reactor, at LHSVs of 4, 6 and 32,conversion also rose dramatically to about 50% when diluent was removedand this increase in conversion was accompanied by a substantialincrease in ethylene selectivity and in some cases, a modest descreasein propylene selectivity. In each case involving the catalyst-containingmicrochannel reactor, at LHSVs of 4, 6 and 32, the hydrocarbonconversion and total yield of olefins improved significantly whendiluent was removed from the system. Table 4 below shows the effect ofC3:O₂ ratio in the microchannel reactor ODH v1

TABLE 4 LHSV = 32, ODH v1* C₃:O₂ 2 1 0.5 C3 Conversion (%) 43.8 48.028.8 Propylene Yield (%) 19.0 11.5 9.8 Olefin Yield (%) 30.6 22.1 10.9CO_(x) Selectivity (%) 19.2 42.7 61.5 CO/CO₂ Ratio 2.78 1.12 0.92 O₂Conversion (%) 65.3 66.4 38.5 Selectivity to CO₂ (%) 5.1 20.2 32.1Selectivity to CO (%) 14.1 22.5 29.4 Selectivity to CH₄ (%) 9.8 10.0 0.7Selectivity to C₂H₆ (%) 1.2 1.2 0.1 Selectivity to Propylene (%) 43.523.9 33.9 Selectivity to Ethylene (%) 26.4 22.2 3.9 Catalyst Bed Temp (°C.) 538 540 544 *was not heat treated or given surface coating prior tooperationAt LHSV=32 there is an optimum at C3:O2=2:1 where both propene yield andtotal olefin yield are maximized and carbon oxides are minimized. Asoxygen in the feed increases both propene yield and total olefin yieldfall off dramatically. This is surprising since it would be expectedthat with more O2 in the feed (C3:O2=0.5) there would be more ODH.However, it is believed that with a pre-mixed feed, there will beundesirable gas phase combustion reactions occurring ahead of thecatalyst that lead to CO2 formation. This reaction consumes more oxygenthan ODH.C3H8+5O2=3CO2+4H2O  (1)VersusC3H8+0.5O2=C3H6+H2O  (2)As the amount of oxygen in the feed decreases, the rate of reaction 1will also fall and leave more time and oxygen available for ODHresulting in lower carbon oxides and higher yields of desired olefinicproducts. In view of these unexpected results, with pre-mixed feeds, wewould expect that as the C3:O2 ratio falls (more O2 in feed), higherlinear velocities (higher LHSVs) would be needed to reduce gas phasereactions occurring before the catalyst bed and optimise conversion touseful products.Table 5 shows the differences between the quartz reactor and themicrochannel reactor ODH v1 at the optimum C3:O2 ratio of 2:1 identifiedin Table 3 above.

TABLE 5 Comparison of LHSV Effects in a Microchannel Pellet ODH v1 and aQuartz Tube C₃:O₂ = 2.:1 Quartz Quartz Device Tube ODHv1* Tube ODHv1*ODHv1* LHSV (v/v//hr) 8 8 32 32 157 C3 Conversion 25.0 37.5 27.4 43.879.9 (%) Propylene 12.1 13.6 12.3 19.0 13.0 Yield (%) Olefin Yield (%)14.0 21.0 14.2 30.6 43.6 CO_(x) Selectivity 41.6 32.8 46.6 19.2 22.1 (%)CO/CO₂ Ratio 1.32 1.17 1.71 2.76 4.55 O₂ Conversion 79.5 78.9 79.4 65.336.9 (%) Selectivity to 17.9 15.1 17.2 5.1 4.0 CO₂ (%) Selectivity to23.7 17.7 29.4 14.1 18.2 CO (%) Selectivity to 2.1 9.9 1.4 9.8 19.7 CH₄(%) Selectivity to 0.2 1.3 0.1 1.2 3.6 C₂H₆ (%) Selectivity to 12.1 13.612.3 19.0 13.0 Propylene (%) Selectivity to 7.7 19.7 6.9 26.4 38.3Ethylene (%) Catalyst Bed 549 538 583 538 512 Temp (° C.) *was not heattreated or given surface coating prior to operationThe table above shows that conversion increases at higher LHSV (shortercontact time). This is again an unexpected result. As stated abovenormally conversion increases with decreasing LHSV (longer contacttime). In the examples in Table 5 we believe that homogeneous gas phasereactions are also occurring ahead of the catalyst zone leading toformation of carbon oxides. These reactions consume more oxygen thanODH.C3H8+5O2=3CO2+4H2O  (1)VersusC3H8+0.5O2=C3H6+H2O  (2)

As LHSV increases, linear velocity increases and residence time in thegas phase ahead of the catalyst decreases. This reduces reaction 1 asseen in the examples producing less carbon oxides and leaving moreoxygen for the ODH reaction resulting in a higher conversion as shown.Going from LHSV=8 to LHSV=32 in the microchannel reactor ODH v1, thetemperature remains constant, selectivity to propylene and ethylene bothincrease substantially while the selectivity to COx decreasessubstantially. This contrasts with the quartz fixed bed where thetemperature rises, olefin selectivity falls and COx rises.

The overall effect is that at LHSV=32 the olefin yield from themicrochannel reactor is more than twice that of the conventionalreactor. The data at LHSV 157 at a higher C3:O2 ratio of 1:1, show thatthe total olefin yield can be increased even higher in the microchannelreactor without having a large increase in temperature which wouldincrease the unselective gas phase reactions and lead to higher yieldsof carbon oxides in a conventional reactor.

To investigate the effect of surface coatings on the ODH reaction testswere run with alumina and silica coated microchannel reactors ODH v3a atvarious LHSVs and temperatures. The results are shown in Table 6.

TABLE 6 Effect of Surface Treatment & LHSV on Conversion and Selectivityin a Microchannel Pellet C₃:O₂ = 2:1, ODH v3a Surface Treatment AluminaSilica Alumina Alumina Silica Alumina Silica Alumina Silica LHSV(v/v//hr) 8 8 8 32 32 157 157 157 157 C3 Conversion (%) 19.6 24.7 14.817.0 24.1 16.1 21.1 40.5 32.0 Propylene Yield (%) 5.9 8.2 2.4 5.2 8.55.5 8.5 14.3 15.2 Olefin Yield (%) 7.6 8.7 2.6 5.7 8.7 5.9 8.7 24.8 17.4CO_(x) Selectivity (%) 58.3 63.8 82.1 65.3 63.4 63.1 58.6 28.0 43.5CO/CO₂ Ratio 0.71 1.21 0.66 0.75 1.18 1.37 1.55 2.11 1.59 O₂ Conversion(%) 90.6 86.8 92.0 85.5 82.2 54.6 50.6 71.1 76.6 Selectivity to CO₂ (%)33.9 28.9 49.6 37.4 29.1 26.6 23.0 9.0 16.8 Selectivity to CO (%) 24.334.9 32.5 27.9 34.3 36.5 35.6 19.0 26.7 Selectivity to CH₄ (%) 2.8 0.70.3 0.7 0.2 0.4 0.2 9.2 2.1 Selectivity to C₂H₆ (%) 0.2 0.1 0.1 0.1 0.10.1 0.1 1.6 0.0 Selectivity to Propylene (%) 5.9 8.2 2.4 5.2 8.5 5.5 8.514.3 15.2 Selectivity to Ethylene (%) 8.8 2.3 1.3 3.1 1.0 1.9 1.1 26.17.1 Catalyst Bed Temp (° C.) 537 539 499 541 542 552 546 599 631Comparing examples in Table 6 at the same LHSV and temperature it can beseen that the silica coating has a beneficial effect compared toalumina. In these non-optimised experiments, the silica coating giveshigher conversions, higher propene selectivity, higher total olefinyields and lower CO2 and methane at all LHSVs.

Example 2 Propane Oxidative Dehydrogenation in Conventional andMicrochannel Reactors

The catalyst in this example is the same that was used in Example 1,except that 5 weight % MgO was added as a binder to the powderedcatalyst. It was tested in the microchannel device ODH v2 (C in Table 7below) and in a comparative test in a 4 mm I.D. quartz tube (B in Table7 below). In the quartz tube, this catalyst demonstrated substantiallylower activity compared to the catalyst of Example 1 which was alsotested in a larger diameter quartz tube reactor (A in Table 7). Theresults show that despite the lower performance of this catalyst in athe smaller diameter quartz tube fixed bed reactor, the catalystout-performs both quartz tube reactors when run in a microchannelreactor ODH v2 at the same temperature as measured in the catalyst bed.Comparing B and C shows that for the same catalyst in the microchannelreactor, conversion is 94% higher, propylene selectivity isapproximately the same and propylene yield is 85% higher. Normally inoxidation reactions the selectivity to the desired product falls asconversion increases. Here the results show that despite a very largeincrease in conversion in the microchannel reactor, selectivity topropene only falls by 1.9%

TABLE 7 Performance of Quartz Tubes vs. the ODH v2 Pellet LHSV = 32,C₃:O₂ = 1:1 B. Quartz A. Quartz Tube C. Tube 4 mm Microchannel 10 mmI.D. I.D. + 5% Device ODH No binder MgO binder v2 C3 Conversion (%) 38.722.7 44.0 Propylene Yield (%) 15.1 10.0 18.5 Olefin Yield (%) 18.6 11.921.9 O₂ Conversion (%) 72.3 36.5 99.9 Selectivity to CO₂ (%) 19.7 23.825.5 Selectivity to CO (%) 31.0 23.0 20.6 Selectivity to CH₄ (%) 1.1 3.53.6 Selectivity to C₂H₆ (%) 0.1 0.1 0.5 Selectivity to Propylene 38.944.0 42.1 (%) Selectivity to Ethylene (%) 9.2 8.3 7.7 Inlet GasTemperature 540° C. 540° C. 623° C. Catalyst Bed Temperature* 636° C.627° C. 628° C. *The temperature in the outlet side of the catalyst bedin the tube A, the tube wall temperature at the bottom of the catalystbed in the tube B or the average web temperature in the ODH v2 pelletSimilar tests were conducted for a C3:O2 feed ratio of 2:1 and resultsare shown in Table 8 below. Test A in Table 8 is for the catalyst fromExample 1 (no MgO binder) tested in a 10 mm I.D. quartz tube. Test B isthe same catalyst tested in the ODH v1 microchannel reactor. Test C isthe catalyst with 5% MgO binder tested in the ODH v2 microreactor. Asshown in Table 5 above, comparing A and B shows that the unboundcatalyst performs better in a microchannel reactor than a quartz tubefixed bed reactor in terms of conversion and olefins yield. This is eventhough the temperature rises in the fixed bed (A) by 43° C. and might beexpexted to increase conversion compared to the microchannel reactorswhere the temperature is uniform along the channel to within 2 or 3degrees. Test C shows that the catalyst that is less active in a fixedbed reactor (according to tests A and B in Table 7 above) achieves asuperior performance in a microchannel reactor ODH v2 operated atapproximately the same catalyst bed temperature as the fixed bed (A).Conversion is increased by 76%, selectivity to carbon oxides is reducedby 33% and olefins yield is increased by 119% over the fixed bed.

TABLE 8 Performance of Quartz Tubes vs. the ODH v2 Pellet LHSV = 32,C₃:O₂ = 2:1 A. Quartz B. C. Tube Microchannel Microchannel 10 mm DeviceDevice I.D ODH v1 ODH v2 No binder No Binder +5% MgO C3 Conversion (%)27.4 43.8 48.4 Propylene Yield (%) 12.3 19.1 19.3 Olefin Yield (%) 14.230.6 31.1 O₂ Conversion (%) 79.4 65.3 98.7 Selectivity to CO₂ (%) 17.25.1 11.5 Selectivity to CO (%) 29.4 14.1 14.7 Selectivity to CH₄ (%) 1.49.8 8.4 Selectivity to C₂H₆ (%) 0.1 1.2 1.0 Selectivity to Propylene44.9 43.5 40.0 (%) Selectivity to Ethylene (%) 6.9 26.4 24.3 Inlet GasTemperature 540° C. 540° C. 580° C. Catalyst Bed Temperature* 583° C.538° C. 577° C. *The temperature in the outlet side of the catalyst bedin the tube A, the tube wall temperature at the bottom of the catalystbed in the ODH v1 or the average web temperature in the ODH v2 pelletIn Table 9a comparison is made between the performance of the quartztube (unbound catalyst) and the microchannel pellet ODH v2 (+5% MgObinder). In this case two differences can be seen between theperformance of the quartz tube and the microchannel device. Comparisonruns were selected such that the average bed temperatures in themicrochannel device were as close as possible to the bed temperature inthe quartz tube. The first difference to note is that the microchanneldevice has its optimal yield of olefins at a C3:O2 ratio of 2:1 whilethe optimal yield for the quartz tube is at a C3:O2 ratio of 1:1. Thelocation of the optimal yield in the microchannel device is alsosupported by data reported in Table 4 above. The second difference isthat the microchannel device is able to produce a higher yields of bothpropylene and total olefins at both C3:O2 of 1:1 and 2:1 by operatingclose to isothermally near the quartz tube catalyst bed temperature.

TABLE 9a C₃:O₂ = 1:1 C₃:O₂ = 2:1 Tube ODH v2 Tube ODH v2 Conversion ofPropane (%) 38.7 44.0 27.4 46.8 Selectivity to Propylene (%) 38.9 42.144.9 39.3 Propylene Yield (%) 15.1 18.5 12.3 18.4 Total OlefinSelectivity (%) 48.1 49.8 51.8 62.5 Olefin Yield (%) 18.6 21.9 14.2 29.3Inlet Gas Temperature 540° C. 623° C. 540° C. 560° C. Catalyst BedTemperature* 636° C. 628° C. 583° C. 595° C.Table 10 shows the influence of temperature on the performance ofpropane ODH conducted in the microchannel reactor ODH v2 using the boundcatalyst at a contact time of 250 milliseconds calculated on total gasflow and a C3:O2 feed ratio of 1:1. The results show that for thecatalyst tested the temperature of the device needs to be elevated above540° C. in order to obtain significant conversion and yields of both C3and C2 olefins. At C3:O2=1:1, the highest yield seen here is at thehighest temperature tested, i.e. 650° C.

TABLE 10 Influence of Temperature in the Microchannel Pellet ODH v2 CT =250 ms, C₃:O₂ = 1:1 Conversion of Propane (%) 22.8 39.2 44.0 50.0Selectivity to Propylene (%) 37.3 42.3 42.1 38.8 Propylene Yield (%) 8.516.6 18.5 19.4 Total Olefin Selectivity (%) 39.0 46.1 49.5 51.8 OlefinYield (%) 8.9 18.1 21.9 25.9 Inlet Gas Temperature 541° C. 597° C. 623°C. 648° C. Catalyst Bed Temperature* 540° C. 602° C. 628° C. 654° C.*The average web temperature in the ODH v2 pelletTable 11 shows similar results for a C3:O2 feed ratio of 2:1. Theresults show that the temperature at which optimal olefins yield isobtained changes with the ratio of C3:O2. Here the highest yield is seenat 600° C.

TABLE 11 Influence of Temperature in the Microchannel Pellet ODH v2 CT =250 ms, C₃:O₂ = 2:1 C3 Conversion (%) 46.8 42.9 40.3 Propylene Yield (%)18.4 17.6 17.7 Olefin Yield (%) 29.3 25.4 23.4 CO_(x) Yield (%) 13.013.7 13.8 O₂ Conversion (%) 99.4 99.9 100.0 Selectivity to CO₂ (%) 13.416.9 18.9 Selectivity to CO (%) 14.3 15.1 15.4 Selectivity to CH₄ (%)8.7 7.9 6.8 Selectivity to C₂H₆ (%) 1.1 0.9 0.9 Selectivity to Propylene(%) 39.3 41.1 44.0 Selectivity to Ethylene (%) 23.2 18.2 14.0 Inlet GasTemperature 596° C. 623° C. 647° C. Catalyst Bed Temperature* 595° C.624° C. 650° C. *The temperature in the outlet side of the catalyst bedin the LCIC tube, the tube wall temperature at the bottom of thecatalyst bed in the Velocys tube or the average web temperature in theODH v2 pelletTable 12 shows similar results for a C3:O2 feed ratio of 2.6:1. Hereagain the highest yield of olefins is seen at 575° C., a lowertemperature than at the feed ratios reported above. It appears that thetemperature at which ‘optimal’ yields are obtained increases withdecreasing C3:O2 ratio.

TABLE 12 Influence of Temperature in the Microchannel Pellet ODH v2 CT =250 ms, C₃:O₂ = 2.6:1 575 625 C3 Conversion (%) 42.2 31.9 PropyleneYield (%) 18.3 15.6 Olefin Yield (%) 28.1 19.1 CO_(x) Yield (%) 10.211.0 O₂ Conversion (%) 98.2 99.9 Selectivity to CO₂ (%) 13.8 19.3Selectivity to CO (%) 10.4 15.3 Selectivity to CH₄ (%) 8.4 5.4Selectivity to C₂H₆ (%) 0.9 0.0 Selectivity to Propylene (%) 43.4 48.9Selectivity to Ethylene (%) 23.2 11.1 Inlet Gas Temperature 578° C. 623°C. Catalyst Bed Temperature* 575° C. 625° C. *The temperature in theoutlet side of the catalyst bed in the LCIC tube, the tube walltemperature at the bottom of the catalyst bed in the Velocys tube or theaverage web temperature in the ODH v2 pelletTable 13 shows the influence of C₃:O₂ feed ratio at constant temperaturein the microchannel pellet ODH v2. When the temperature of themicrochannel is held constant and the C3:O2 ratio is changed it can beseen that the propylene yield falls with increasing C3:O2 ratio but thatthe total olefin yield appears to pass through a maximum in the regionof 2:1.

TABLE 13 Influence of C₃:O₂ Ratio at Constant Temperature in theMicrochannel Pellet ODH v2 LHSV = 32 C₃:O₂ = 1:1 C₃:O₂ = 2:1 C₃:O₂ =2.6:1 Conversion of Propane (%) 44.0 42.9 31.9 Selectivity to Propylene42.1 41.1 48.9 (%) Propylene Yield (%) 18.5 17.6 15.6 Total OlefinSelectivity (%) 49.8 59.3 60.0 Olefin Yield (%) 21.9 25.4 19.1 Inlet GasTemperature 623° C. 623° C. 623° C. Catalyst Bed Temperature* 628° C.624° C. 625° C. *The average web temperature in the ODH v2 pelletTable 14 below, shows the influence of contact time in the microchannelpellet ODH v2 at constant temperature. As in Table 5. above, the resultsalso show a surprising trend in that conversion increases as contacttime over the catalyst decreases. This is opposite to what might beexpected in conventional systems where increasing contact time usuallyresults in an increased conversion. Here again we believe thathomogeneous gas phase reactions are also occurring ahead of the catalystzone leading to formation of carbon oxides. These reactions consume moreoxygen than ODH.C3H8+5O2=3CO2+4H2O  (1)VersusC3H8+0.5O2=C3H6+H2O  (2)As contact time decreases, linear velocity increases and residence timein the gas phase ahead of the catalyst decreases. This reduces reaction1 as seen in the examples producing less carbon oxides and leaving moreoxygen for the ODH reaction resulting in a higher conversion, lower COxand higher olefin yields as shown.

TABLE 14 Influence of Contact Time in the Microchannel Pellet ODH v2C₃:O₂ = 2.0:1 Contact Time (ms) 61 82 122 250 C3 Conversion (%) 57.352.7 47.8 46.8 Precat. Cntct Time (ms) 26 35 53 105 Propylene Yield (%)18.2 19.2 18.4 18.4 Olefin Yield (%) 37.8 33.1 32.1 29.3 CO_(x) Yield(%) 11.0 11.5 10.1 13.0 O₂ Conversion (%) 96.9 97.7 99.8 99.4Selectivity to CO₂ (%) 5.6 7.0 8.3 13.4 Selectivity to CO (%) 13.6 14.812.8 14.3 Selectivity to CH₄ (%) 12.6 13.1 10.2 8.7 Selectivity to C₂H₆(%) 2.2 2.2 1.6 1.1 Selectivity to Propylene (%) 31.7 36.4 38.6 39.3Selectivity to Ethylene (%) 34.3 26.5 28.5 23.2 Inlet Gas Temperature597° C. 597° C. 597° C. 596° C. Catalyst Bed Temperature* 603° C. 602°C. 599° C. 595° C. *The temperature in the outlet side of the catalystbed in the LCIC tube, the tube wall temperature at the bottom of thecatalyst bed in the Velocys tube or the average web temperature in theODH v2 pellet

Example 3 Improved Catalyst Composition Containing Mo

Mg_(4.5) Mo₁V₁O_(n) (or 43.8 wt % MgO; 21.7 wt % V₂O₅ and 34.4 wt %MoO₃; SA 31 m²/g) and Mg_(7.75)Mo_(0.1)V₁O_(n) (or 74.8 wt % MgO; 21.7wt % V₂O₅ and 3.5 wt % MoO₃) were prepared and tested in the quartzreactor (LHSV=32, C₃:O₂=1:1). The results for Mo:V=0.1:1 show similarresults as the binary oxides with slightly increased selectivity topropene at the expense of CO. Mo:V=1:1 catalyst gives significantlyhigher propane conversion and comparable propene selectivity, meaninghigher overall yield than Mg—V—O. CO_(X) selectivity is nearly halved,whereas methane and especially ethylene are significantly increased sothat the total olefin yield is doubled. This leads to a much lowerexotherm than the conventional catalyst.

Mg—V—O vs. Mg—Mo—V—O ODHv3a ODHv3a Quartz Tube Quartz Tube Quartz TubeSilica Coated Silica Coated Catalyst Mg—V—O Mg—Mo—V—O Mg—Mo—V—OMg—Mo—V—O Mg—Mo—V—O Mo:V Ratio N/A 0.1:1 1:1 1:1 1:1 C₃:O₂ 1 1 1 1 2 C3Conversion (%) 38.7 38.0 65.9 16.2 10.8 Propylene Yield (%) 15.0 16.123.0 2.4 3.6 Olefin Yield (%) 18.6 19.9 40.9 2.8 3.9 CO_(x) Selectivity(%) 50.7 46.4 27.1 82.3 62.6 CO/CO₂ Ratio 1.6 1.2 2.5 2.1 1.6 O₂Conversion (%) 72.3 63.1 58.7 25.2 31.2 Selectivity to CO₂ (%) 19.7 20.97.8 26.1 24.4 Selectivity to CO (%) 31.0 25.5 19.3 56.2 38.2 Selectivityto CH₄ (%) 1.1 1.2 8.8 0.2 0.8 Selectivity to C₂H₆ (%) 0.1 0.1 2.1 0.10.1 Selectivity to Propylene (%) 38.9 42.4 34.9 16.1 32.9 Selectivity toEthylene (%) 9.2 9.9 27.1 1.3 3.6 Catalyst Bed Temp (° C.) 636 558 583537 537

Example 4 Ethane Oxidative Dehydrogenation in Conventional andMicrochannel Reactors

The ODH catalyst used in the testing for this example was aSm₂O₃—Li—Cl/MgO catalyst containing 5.2 wt % Sm₂O₃, 3.4 wt % Li, and12.5 wt % Cl supported on Mgo. The powder catayst was pelletized,crushed and sieved into the size range 150-210 μm. The powder form wasfound to have a surface area of 21 m²/g. Felt based catalysts wereprepared by ball milling the native mixed oxide the slurry coating on tothe FeCrAlY substrate. The surface area of the felt was found to be 81m²/g. Prior to use the catalyst was heated in are to 200° C. at 1°C./min and then heated under air to the maximum reation temperature at arate of 5° C./min. All testing done on ethane ODH was performed usingair as the source of oxygen.

Studies of ethane ODH employing the catalysts described above wereperformed in various devices, quartz tubes (ID 4 mm), Inconel 625 tubes(ID 4.8 mm) and microchannel pellets ODH v1 and ODH v3b underexperimental conditions ranging from C2:O2 ratios of 2:1 to 10.1:1,contact times of 1016 ms to 20 ms and temperatures from 500° C. and 800°C. Illustrative comparisons between the various devices are made inTables 16, 17 and 18.

A comparison between an blank (containing no catalyst) quartz tube, aquartz tube with powdered catalyst, an Inconel 625 tube with powderedcatalyst and an ODH v3b microchannel test pellet containing a felt canbe found in Table 16. It can be seen that the inclusion of catalystincrease the conversion of ethane and selectivity to ethylene in bothquartz tube (17.0% and 13.0% respectively) and the in the microchannelreactor (19.5% abd 15.0% respectively) as compared to the blank (12.5%and 10.9% respectively). The importance of surface passivation can beseen when the results of the quartz tube, Inconel 625 tube and the ODHv3b device are compared. The results for the untreated Inconel 625 tubeindicate that the surface promotes combustion resulting in a 67.4%selectivity to COx as compared to 15.5% for the quartz blank and 16.0%for the ODH v3b device. In addition the overall combustion is lower inthe Inconel 625 tube, 8.3% as compared to 17.0% in the quartz tube and19.5% in the ODH v3b device. As was noted in example 1 in the case ofpropane ODH deep combustion competes with the ODH reaction reducing theoverall conversion of ethane. This is further supported by theobservation that the microchannel has a greater conversion of ethane andyield of ethylene (19.5% and 15.0%) where the CO/CO₂ ratio is 1.46 thanthe quartz tube (17.0% and 13.0%) where the CO/CO₂ ratio is 0.42(indicating that less oxygen was available for ODH).

TABLE 16 Ethane ODH Performance of Quartz Tubes vs. the ODH v3b PelletCT = 250 ms, C₂:O₂ = 10:1, Sm₂O₃ Catalyst, Air as Oxidant Quartz IN625Blank Tube Tube Microchannel Quartz 4 mm 4.8 mm Device Tube ID ID ODHv3b** Catalyst Type None Powder Powder Felt C2 Conversion (%) 12.5 178.3 19.5 Ethylene Yield (%) 10.9 13.0 2.6 15.0 Olefin Yield (%) 10.913.2 2.6 15.0 O₂ Conversion (%) 0.89 100.0 100.0 99.4 Selectivity to CO₂(%) 2.2 10.9 31.0 6.5 Selectivity to CO (%) 6.6 4.6 36.4 9.5 Selectivityto CH₄ (%) 4.1 5.8 1.6 7.1 Selectivity to Propane (%) 0.3 1.1 0.0 0.0Selectivity to Propylene 0.0 1.1 0.0 0.0 (%) Selectivity to Ethylene (%)86.8 76.6 31.0 76.9 Inlet Gas Temperature 650° C. 648° C. 658° C. 652°C. Catalyst Bed Temperature* 650° C. 645° C. 650° C. 647° C. *The tubewall temperature at the bottom of the catalyst bed in the tube or theaverage external temperature over the catalyst bed in the ODH v3b pellet**pellet was sol coated with alumina

The sensitivity of the performance of ethane ODH conducted in Inconel625 tubes and ODH v3b microchannel test pellets is demonstrated by datapresented in Table 17. It can be seen from the data in Table 17 that atapproximately 600° C. the Inconel 625 tubes performed better that theODH v3b microchannel device in terms of conversion of ethane andselectivity to ethylene but at approximately 645° C. the ODH v3bmicrochannel device performed much better that the Inconel 625 tube.

TABLE 17 Ethane ODH IN625 Tube vs. Sol Coated DH Pellet CT = 250 ms,C₂:O₂ = 2.5:1, Sm₂O₃ Catalyst on Felt, Air as Oxidant MicrochannelMicrochannel Device IN617 Device IN625 Tube DH Pellet** IN625 Tube ODHv3b** Catalyst Type Powder Felt Powder Felt C2 Conversion (%) 16.7 11.523.5 29.5 Ethylene Yield (%) 5.1 2.0 0.4 12.0 Olefin Yield (%) 5.1 2.00.4 12 O₂ Conversion (%) 99.9 64.6 100.0 99.0 Selectivity to CO₂ (%)65.5 76.0 36.3 38.8 Selectivity to CO (%) 3.7 5.7 61.6 16.2 Selectivityto CH₄ (%) 0.2 1.2 0.6 4.3 Selectivity to Propane (%) 0.0 0.0 0.0 0.0Selectivity to Propylene (%) 0.0 0.0 0.0 0.0 Selectivity to Ethylene (%)30.6 17.1 1.5 40.7 Inlet Gas Temperature 608° C. 601° C. 657° C. 649° C.Catalyst Bed Temperature* 589° C. 600° C. 641° C. 645° C. *Thetemperature in the outlet side of the catalyst bed in the LCIC tube, thetube wall temperature at the bottom of the catalyst bed in the Velocystube or the average external temperature over the catalyst bed in theODHv3bpellet **pellet was sol coated with alumina

It was desired to not only demonstrate the inherent ability of amicrochannel to be operated in a close to isothermal manner and therebyimprove selectivity in ODH but to demonstrate as well the ability toapply active cooling adjacent to the channel containing the catalyst andthereby maintain more closely control the reaction conditions. This wasattempted in a ODH v1 pellet by operating with and without coolant airflowing in the cooling channels, the results can be seen in Table 18.

From the data it appears that when columns 2 and 3 of Table 18 arecompared that the application of active cooling did little to influencethe outcome of the ODH reaction. In addition to what is reported heredata was collected under several conditions with coolant air flowing at2200 and 5900 sccm and in all cases no significant influence was notedon the selectivities and conversion. The lack of influence of theapplication of cooling on the reaction was later traced a fabricationerror made in this series of test pellets in which the coolantdistribution feature was completely blocked when the coolant channelswere sealed. The blocking lead to a situation in which the bulk of thecoolant would by-pass the portions of the coolant channel adjacent tothe catalyst.

What is apparent from Table 18 is that when a microchannel device isoperated close to the measured peak temperature of a quartz tube boththe conversion of ethane and selectivity to ethylene are increased andthis increase is by approximately the same factor, conversion beingincreased by 1.87 times and selectivity by 1.84 times.

TABLE 18 Ethane ODH Quartz tube vs. ODH v1 (alumina sol coated) with andwithout Cooling CT = 20 ms, C₂:O₂ = 10:1, Sm₂O₃ Catalyst, Air as OxidantODH v1 ODH v1 1 Felt, 1 Felt, Quartz Tube No Cooling 2200 SCCM C2Conversion (%) 7.7 14.4 14.7 Ethylene Yield (%) 5.9 10.9 11.2 OlefinYield (%) 6.0 10.9 11.2 O₂ Conversion (%) 75.1 N/A N/A Selectivity toCO₂ (%) 16.2 6.7 6.7 Selectivity to CO (%) 1.4 11.1 10.7 Selectivity toCH₄ (%) 3.5 5.5 5.4 Selectivity to Propane (%) 1.6 1.2 1.1 Selectivityto Propylene (%) 1.3 0.0 0.0 Selectivity to Ethylene (%) 76.1 75.5 76.1Inlet Gas Temperature 618° C. 669° C. 665° C. Catalyst Bed Temperature*678° C. 671° C. 668° C. *The tube wall temperature at the bottom of thecatalyst bed in the tube or the average external temperature over thecatalyst bed in the ODH v3b pellet

1. Apparatus for oxidatively dehydrogenating a hydrocarbon, comprising:a microchannel reaction chamber; and an oxidative dehydrogenationcatalyst disposed in the microchannel reaction chamber; a microchannelmixer that is integral with the microchannel reaction chamber; andcomprising: an oxygen channel adjacent to said microchannel reactionchamber and separated by an oxygen channel wall, wherein aperturesthrough said oxygen channel wall form passageways between the oxygenchannel and the reaction chamber.
 2. The apparatus of claim 1 whereinthe microchannel reaction chamber comprises microchannel grooves. 3.Apparatus for oxidatively dehydrogenating a hydrocarbon, comprising: amicrochannel reaction chamber; and an oxidative dehydrogenation catalystcomprising an oxide or phosphate of a metal selected from the groupconsisting of Li, Mo, V, Nb, Sb, Sn, Zr, Cr, Mg, Mn, Ni, Co, Ce,rare-earth metals, and mixtures thereof; and wherein the oxidativedehydrogenation catalyst is disposed in the microchannel reactionchamber and comprises one of the following forms: a) a particulatecatalyst; or b) a porous insert; or c) a catalyst wall coatingcomprising a first layer formed between a reaction chamber wall and asecond layer; wherein the reaction chamber wall, first layer and secondlayer have different compositions, wherein the first layer has athickness of at least 0.1 micrometers.
 4. The apparatus of claim 3wherein the oxidative dehydrogenation catalyst fills a cross-sectionalarea of the microchannel so that there is no bulk flow path through themicrochannel.
 5. The apparatus of claim 3 wherein a silica coating isdisposed on the wall.
 6. The apparatus of claim 3 wherein the oxidativedehydrogenation catalyst comprises a felt.
 7. The apparatus of claim 3further comprising Cu, Ag, or Sn.
 8. The apparatus of claim 3 whereinthe catalyst comprises Mg and V.
 9. The apparatus of claim 3 comprisingMg, V, and Mo, wherein the molar ratio of Mo:V is in the range of 0.5 to2.
 10. A catalytic system for oxidatively dehydrogenating a hydrocarbon,comprising: a reaction chamber; and an oxidative dehydrogenationcatalyst disposed in the reaction chamber; wherein the system ischaracterizable by a catalytic activity such that when propane and O₂,with no diluents, in a 1:1 ratio are fed into the reaction chamber at anLHSV of 32 and a catalyst temperature of 580° C., there is a propaneconversion of at least 30% and an olefin yield of at least 20%.
 11. Thecatalytic system of claim 10 wherein the system is characterizable by acatalytic activity such that when propane and O₂, with no diluents, in a1:1 ratio are fed into the reaction chamber at an LHSV of 32 and acatalyst temperature of 580° C., there is a propane conversion of atleast 30 to about 50% and an olefin yield of at least 20 to 31%.
 12. Thecatalytic system of claim 10 wherein the system is characterizable by acatalytic activity such that when the catalyst is replaced by an equalvolume of a catalyst consisting of a Mg—V—O catalyst containing 81 wt %MgO and 19 wt % V₂O₅ with surface area of about 100 m²/g pelleted to asize between 250-400 μm and then pre-treated prior to reaction at 500°C. in 40ml/min O₂ for 1 h is inserted into the reaction chamber, andthen propane and O₂, with no diluents, in a 1:1 ratio are fed into thereaction chamber at an LHSV of 32 and a catalyst temperature of 580 C,there is a propane conversion of at least 30% and an olefin yield of atleast 20%.
 13. The catalytic system of claim 10 wherein the oxidativedehydrogenation catalyst disposed in the microchannel reaction chambercomprises one of the following forms: d) a particulate catalyst; or e) aporous insert; or f) a catalyst wall coating comprising a first layerformed between a reaction chamber wall and a second layer; wherein thereaction chamber wall, first layer and second layer have differentcompositions, wherein the first layer has a thickness of at least 0.1micrometers.
 14. Apparatus for oxidatively dehydrogenating ahydrocarbon, comprising: a process channel comprising a u-turn; whereinthe u-turn is disposed between a forward section of the process channeland a return section of the process channel; a microchannel reactionchamber disposed in the process channel; an oxidative dehydrogenationcatalyst disposed in the microchannel reaction chamber; wherein thereturn channel comprises at least one alkene and/or aralkene; and anoxygen channel adjacent to said process channel and separated by anoxygen channel wall, wherein apertures through said oxygen channel wallform passageways between the oxygen channel and the process channel. 15.The apparatus of claim 14 wherein the u-turn is disposed between aforward section of the process channel and a return section of theprocess channel; and wherein the oxygen channel is disposed between theforward section and the return section.
 16. The apparatus of claim 14further comprising a second process channel comprising a second u-turnand a second forward section and a return section, wherein the u-turn isdisposed between the second u-turn and the return section, and whereinthe process channel and the second process channel share a common returnchannel.
 17. The apparatus of claim 16 wherein the oxygen channel isdisposed between the forward section and the return section.
 18. Theapparatus of claim 14 wherein the return section comprises an outlet;wherein the microchannel reaction chamber is disposed in an area of thereturn section that is closer to the u-turn than to the outlet.
 19. Theapparatus of claim 14 further comprising a heat transfer channeladjacent to the forward channel.
 20. The apparatus of claim 19 whereinthe heat transfer channel comprises an endothermic reaction catalyst.21. The apparatus of claim 14 wherein said apertures through said oxygenchannel wall form passageways between the oxygen channel and themicrochannel reaction chamber.
 22. The apparatus of claim 14 whereinwalls of the process channel comprise a silica coating.
 23. Theapparatus of claim 14 comprising a microchannel mixer that is integralwith the microchannel reaction chamber.